Systems and methods for separating alkane gases with applications to raw natural gas processing and flare gas capture

ABSTRACT

The present invention is a field-deployable system for separating methane and natural gas liquids (NGLs) from a raw gas stream comprising a compressor; a dehydrator; a refrigerator having one or more stages; and a separation subsystem adapted to separate the raw gas stream into three product streams including a methane stream that is at least 80% methane, an ethane-rich stream, and a NGLs stream having a vapor pressure of no more than 250 psia at 38° C. The methane stream is sufficiently lean to be useable in existing natural gas engines without modification. The NGLs stream has a sufficiently low vapor pressure to be transportable in standard propane containers. The ethane-rich stream may be utilized within the system itself to power its own operations. The system can be utilized to reduce flaring from liquids-rich gas production sites to an absolute minimum, produce natural gas liquids for transport, and provide dry methane gas suitable for use in portable field generators.

REFERENCE TO RELATED APPLICATIONS

This application is a non-provisional of and claims priority from U.S. Ser. No. 61/836,220, filed on Jun. 18, 2013, entitled “Mobile Alkane Gas Separator,” the entirety of which is hereby incorporated by reference herein.

FIELD OF THE INVENTION

The present invention relates to enabling the utilization of raw natural gas, such as flare gas, for power generation and liquids capture. More specifically, this invention relates to a mobile system for separating raw natural gas into a high-quality methane gas stream, an ethane-rich gas stream, and a natural gas liquids stream having a low vapor pressure.

BACKGROUND OF THE INVENTION

The statements in this section merely provide background information related to the present disclosure and may not constitute prior art.

Currently liquids-rich raw natural gas is being flared in large quantities at numerous locations by oil producers. This activity entails significant loss of income that could be earned by selling the flared natural gas liquids. Still more financial losses are entailed by failing to make use of the methane content of the flared gas to generate power. As a result, such oil producers have to buy their electric power from the grid, or even worse, generate it themselves at significant cost (typically $0.40/kWh) through the use of on-site diesel generators consuming expensive fuel. Furthermore, the large-scale flaring of natural gas has raised environmental issues that could cause state and/or federal regulators to take action to fine, shutdown, or highly regulate their operations.

The United States oil and gas industry annually flared approximately 7.1 billion cubic meters (bcm), or 250 billion cubic feet (bcf) in 2011 (Source: Global Gas Flaring Reduction Partnership, Estimated Flared Volumes from Satellite Data, 2007-2011, 2013.) It is estimated that in the U.S. state of North Dakota alone, where oil production has increased to 534,000 barrels per day by the end of 2011, around 30% of North Dakota's produced associated gas is flared, nearly 190 million cubic feet per day (190,000 mcf/day) (Source: Wocken, C. A.; Stevens, B. G.; Almlie, J. C.; Schlasner, S. M., End-Use Technology Study—An Assessment of Alternative Uses for Associated Gas, National Energy Technology Laboratory, Pittsburgh, Pa., April 2013.) This is estimated to produce harmful CO₂ emissions of 3.6 million tons (Mt) CO₂-equivalents per year while producing no useful product. Canada also has significant flare gas resources. It is estimated that Canada flares 2.4 billion m³ per year (Source: Global Gas Flaring Reduction Partnership, Estimated Flared Volumes from Satellite Data, 2007-2011, 2013.) It is estimated that the Canadian province of Alberta alone flares 868 million m³ and vents another 333 million m³, producing harmful CO₂ emissions of 17 Mt CO₂-equivalents in Alberta while producing no useful product. (Source: Bott, R. D., Flaring Questions and Answers, 2nd ed., Canadian Centre for Energy Information, 2007.) If this waste flare gas could be utilized for powering drilling rigs and other oil field equipment, significant environmental and economic benefits would accrue.

Furthermore, shale gas is often rich in natural gas liquids (NGLs), especially in North Dakota (Bakken formation). Although natural gas prices are at a historic low, the high concentration of NGLs justifies gathering and processing the gas. Typical Bakken gas may contain as much as 10-12 gallons of NGLs per 1 mcf of wellhead gas. (Source: Wocken, C. A.; Stevens, B. G.; Almlie, J. C.; Schlasner, S. M., End-Use Technology Study—An Assessment of Alternative Uses for Associated Gas, National Energy Technology Laboratory, Pittsburgh, Pa., April 2013.) However, as gas gathering infrastructure is put into place, and statutory restrictions kick in after 12 months after initial drilling, flaring on old wells may be reduced. Unfortunately, newly drilled wells as well as wells far from pipeline infrastructure will continue to be flared.

It is highly financially and environmentally disadvantageous to flare valuable natural gas liquids that could be sold at great profit. It is even more financially and environmentally disadvantageous to spend large amounts of money on diesel fuel for power generation and at the same time flare methane that could be doing the same job. The problem is that liquids-rich raw natural gas cannot be used in generators and cannot be transported by truck.

Therefore, there exists an important need for a solution to address the problem of utilizing raw natural gas to the maximum extent and to minimize or eliminate flaring completely.

Accordingly, as recognized by the present inventors, what are needed are a novel method, apparatus, and system for separating raw natural gas into a natural gas liquids stream that can be easily transported by truck, and a lean methane stream that can be utilized for power generation or other purposes in existing equipment. As recognized by the present inventors, what is also needed is a natural gas separator that is compact, portable, and modular, and which can be easily and quickly delivered to flare gas sites as flaring volumes change and as natural gas infrastructure matures.

Therefore, it would be an advancement in the state of the art to provide an apparatus, system, and method for cost-effectively separating natural gas liquids and lean methane from a raw natural gas source at or near an oil or gas site that flares its associated gas. It would also be an advancement in the state of the art to provide a compact, portable, and modular natural gas separator.

It is against this background that various embodiments of the present invention were developed.

BRIEF SUMMARY OF THE INVENTION

Whereas the problem of flare gas was previously recognized, existing systems were optimized towards one of two approaches. Some existing solutions were designed for extracting NGLs from flare gas, and then flared the remaining methane and ethane because it was unusable in existing gensets. Other existing solutions generated high-quality (lean) methane for CNG or pipelines, and flared the remaining NGLs because it had high ethane content, and hence a high vapor pressure, too high to be transportable in existing propane tanks.

The system design presented in the present application accomplishes both goals simultaneously, an accomplishment that was not possible with the prior art systems. The presently disclosed invention produces both lean methane and simultaneously produces useable and transportable NGLs with a vapor pressure low enough to be transportable in propane tanks (Y-grade).

The inventors have realized as a result of substantial experimental work that an optimal temperature and pressure range for the entire process can be achieved with a unique refrigeration design and unique overall system design. The result is a system that can achieve high-quality separation in a field environment using a mobile processing plant. By utilizing a unique cascade/autocascade refrigeration system to achieve an optimal temperature range, the inventors have been able to reduce the operating pressure of the entire system, as well as reduce the cost and complexity of the entire system. The result is a portable natural gas processing plant that can simultaneously achieve both lean methane separation as well as NGLs capture meeting Y-grade specification.

Accordingly, one embodiment of the present invention is a system for separating methane and natural gas liquids from a raw natural gas stream, comprising a chassis or skid adapted to hold the system for field deployment; a compressor for compressing the raw natural gas stream; a dehydrator for removing water from the raw natural gas stream; a refrigerator for lowering a temperature of the natural gas stream to an optimal temperature range, preferably approximately −40° C. to −70° C., and even more preferably −50° C. to −60° C., and possibly −20° C. to −100° C., depending upon the flare gas composition; and a separation subsystem adapted to separate the natural gas stream into three product streams consisting essentially of a methane stream of at least 80% methane (and preferably at least 85% methane, and more preferably at least 90% methane), an ethane-rich stream comprising typically 20-60% ethane, methane, and possibly other residual hydrocarbons, and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C./100° F. (and preferably no more than 225 psia, and even more preferably no more than 200 psia). The methane stream is of sufficient purity and sufficiently lean to be useable in existing natural gas engines without modifications. The natural gas liquids stream is of sufficiently low vapor pressure to be transportable in standard propane pressure vessels (Y-grade). One common definition for Y-grade is a gas mixture having a vapor pressure of no more than 250 psia at 38° C./100° F. In one preferred embodiment of the present invention, the ethane-rich stream is utilized within the system itself to power operations.

Yet another embodiment of the present invention is the system described above, further comprising a power generator tuned to run on ethane-rich gas adapted to provide the electricity to run the compressor, the refrigerator, and other system components, utilizing the ethane-rich stream as its energy source.

Yet another embodiment of the present invention is the system described above, wherein the compressor compresses the raw natural gas stream to a pressure of no more than approximately 300 psig.

Yet another embodiment of the present invention is the system described above, wherein the chassis is mounted on a trailer having one or more wheels.

Yet another embodiment of the present invention is the system described above, wherein the refrigerator further comprises a high-stage refrigeration loop having at least one heat exchanger for lowering a temperature of the dehydrated flare gas; and a low-stage refrigeration loop having at least one heat exchanger for further lowering the temperature of the dehydrated flare gas.

Yet another embodiment of the present invention is the system described above, wherein the low-stage refrigeration loop is an autocascade loop having mixed refrigerants. Yet another embodiment of the present invention is the system described above, wherein the mixed refrigerants are hydrocarbons.

Yet another embodiment of the present invention is the system described above, wherein the dehydrator employs desiccant beds, preferably zeolite beds. Yet another embodiment of the present invention is the system described above, wherein two desiccant beds are employed in alternation, wherein heat required to dry the two beds is derived from waste heat from a power generator that drives the compressor and the refrigerator.

Yet another embodiment of the present invention is the system described above, wherein the separation subsystem comprises a stripping column. Yet another embodiment of the present invention is the system described above, wherein the separation subsystem comprises a distillation column. Yet another embodiment of the present invention is the system described above, wherein the separation subsystem comprises one or more flash tanks.

Yet another embodiment of the present invention is the system described above, wherein the refrigerator cools the natural gas stream to a temperature range of −40° C. to −70° C.

Yet another embodiment of the present invention is the system described above, wherein the separation subsystem further comprises one or more cyclones to separate liquids from gasses.

Another embodiment of the present invention is a method for reducing flaring, comprising the following steps: (1) bringing a mobile alkane gas separator system to an oil field flaring associated gas; (2) compressing a raw associated natural gas stream utilizing a compressor; (3) removing water from the raw natural gas stream utilizing a dehydrator; (4) lowering a temperature of the natural gas stream utilizing a refrigerator loop; and (5) separating the natural gas stream into three product streams consisting essentially of a methane stream of at least 80% methane (and preferably at least 85% methane, and more preferably at least 90% methane), an ethane-rich stream comprising ethane (typically 20-60% ethane), methane, and possibly other residual hydrocarbons, and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C./100° F. (and preferably no more than 225 psia, and even more preferably no more than 200 psia). The methane stream is of sufficient purity and sufficiently lean to be useable in existing natural gas engines without modifications. The natural gas liquids stream is of sufficiently low vapor pressure to be transportable in standard propane pressure vessels (Y-grade). In one preferred embodiment of the present invention, the ethane-rich stream is utilized within the system itself to power operations.

Yet another embodiment of the present invention is the method described above, further comprising a generation step for generating power by utilizing a power generator tuned to run on ethane-rich gas to provide electricity to run the alkane gas separator, utilizing the ethane-rich stream as its energy source.

Yet another embodiment of the present invention is the method described above, wherein the refrigeration step utilizes an autocascade refrigerator having mixed hydrocarbon refrigerants. Yet another embodiment of the present invention is the system described above, wherein the refrigeration step cools the natural gas stream to a temperature range of −40° C. to −70° C. Yet another embodiment of the present invention includes two, three, or more stages of refrigeration. Yet another embodiment of the present invention is the method described above, wherein the dehydration step employs desiccant beds. Yet another embodiment of the present invention is the system described above, wherein two desiccant beds are employed in alternation, wherein heat required to dry the two beds is derived from waste heat from a power generator. Yet another embodiment of the present invention includes the method described above, wherein the desiccant beds are zeolite beds.

Other features, utilities and advantages of the various embodiments of the invention will be apparent from the following more particular description of embodiments of the present invention.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention will be understood by the following detailed description in conjunction with the accompanying drawings, wherein like reference numerals designate like structural elements, and in which:

FIG. 1 shows a block diagram of one embodiment of a mobile alkane gas separator (MAGS).

FIG. 2 shows a block diagram of one embodiment of a cascade/autocascade refrigeration cycle useable in one embodiment of the MAGS according to another embodiment of the present invention.

FIG. 3 shows a flowchart of a process for separating alkane gases according to one embodiment of the present invention.

FIG. 4 shows a perspective view of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention.

FIG. 5 shows another perspective view of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention.

FIG. 6 shows yet another perspective view of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention.

FIGS. 7 a and 7 b show a detailed piping and instrumentation diagram of the mobile alkane gas separator (MAGS).

FIG. 8 illustrates an example of a use case of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is used for electricity production.

FIG. 9 illustrates another example of a use case of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is compressed into CNG, and used in vehicle fuel or transported to a remote location to power remote gensets.

FIG. 10 illustrates yet another example of a use case of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is converted to liquid fuels for easy transport to a remote site.

FIG. 11 illustrates yet another example of a use case of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is used to create an artificial gas cap or for enhanced oil recovery.

FIG. 12 illustrates yet another example of a use case of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is converted to CO₂ and H₂, which are then used for enhanced oil recovery and electricity production, respectively.

FIG. 13 shows a schematic diagram of one embodiment of a dehydration subsystem useable in one embodiment of the MAGS according to another embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The following description is merely exemplary in nature and is in no way intended to limit the scope of the present disclosure, application, or uses.

DEFINITIONS

The following terms of art shall have the below ascribed meanings throughout this specification, unless otherwise stated.

Throughout this disclosure “MAGS-200,” “full scale” apparatus, or any reference to a single full-scale module, will refer to an apparatus module that can process 200 mcf (thousand cubic feet) of raw natural gas per day. Assuming a typical Bakken gas composition illustrated in Table 1, the MAGS-200 unit would produce approximately 1700 gallons of natural gas liquids (˜46 mcf of gas equivalent, ˜23% of the total volume of flare gas), ˜100-120 mcf of lean methane (˜58% of the total volume of flare gas), and ˜25-40 mcf of an ethane-rich stream (˜19% of the total volume of flare gas). Multiple modules can be combined for higher gas flow rates. These product flow estimates are based on a sample assay of a sample flare gas provided from a sample well by a major Bakken oil company, and are provided for explanation purposes only, and are not intended to be limiting the scope of the present invention in ay way. Different input raw gas compositions would produce different quantities of products.

The symbols mcf, MCF, and kcf will all stand for 1 thousand standard cubic feet of gas. The symbols MMCF, MMcf, and mmcf will all stand for 1 million standard cubic feet, or 1,000 mcf. The word “day” shall mean “a day of operations,” which shall be a 24-hour day, but could also be an 8-hour day, a 12-hour day, or some other amount of operational time.

Natural gas at the wellhead is commonly a mixture of methane (C1) with other hydrocarbons, including ethane (C2), propane (C3), butane (C4), pentane (C5), hexane (C6), and higher (C6+). Wellhead natural gas also contains other compounds such as water vapor (H₂O), hydrogen sulfide (H₂S), carbon dioxide (CO₂), oxygen (O₂), and nitrogen (N₂).

Associated gas is natural gas produced as a by-product of oil drilling, either conventional or unconventional extraction (such as hydraulic fracking for tight oil).

Flare gas is natural gas, usually associated gas, that is flared (burned for no useful purpose) because natural gas pipelines are not in place when the oil well is drilled.

Stranded gas is natural gas, usually associated gas that is flared, that cannot be brought to market either because it is off-shore or too far from natural gas pipelines/infrastructure.

Wet gas is natural gas that contains a high proportion of C2+ components (typically more than 10%). Wet gas is frequently also saturated with water vapor. This is an approximate definition often used by those skilled in the art.

Dry gas is natural gas with typically less than 5% C3+ components, or typically less than 10% C2+ components. This is an approximate definition often used by those skilled in the art.

Natural gas liquids (NGLs) are C3+ components, including propane and heavier hydrocarbons, and may include small amounts of methane and ethane. Other definitions sometimes include ethane as a NGL (natural gas liquid).

Y-grade is an informal standard used to specify NGL requirements. One common definition for Y-grade is a hydrocarbon mixture having essentially no methane and a low ethane content, typically having a vapor pressure of no more than 250 psia at 38° C./100° F. Y-grade is not necessarily limited to 250 psi, but refers to a hydrocarbon mixture that has essentially no methane and a low fraction of ethane, and accordingly a vapor pressure that is lower than raw wellhead gas. Each fractionator that purchases NGLs has different Y-grade requirements. In some embodiments, Y-grade is 150-250 psig vapor pressure. In others, Y-grade may be as high as 600 psig vapor pressure. Standard propane tanks are typically rated for pressures of approximately 200-250 psig at 38° C., so depending on context, Y-grade may refer to a vapor pressure of a mixture that is safe to transport in propane tanks.

LPG is an acronym for Liquefied Petroleum Gas, which is generally a term for compressed, processed gas mixtures of C3+ components, most commonly primarily propane and butane.

CNG is an acronym for Compressed Natural Gas, which is natural gas compressed to a pressure above approximately 2,000 psig, although higher or lower pressures are also possible.

LNG is an acronym for Liquefied Natural Gas, which is natural gas at a pressure and a temperature in which it is a liquid phase.

Joule-Thomson Effect describes the temperature change of a fluid when it passes through a pressure valve or orifice while kept insulated from the environment so that no heat is exchanged. At room temperature, all gases (except hydrogen, helium, and neon) cool upon expansion by the Joule-Thomson effect.

A-gas or A-stream refers to the lean methane stream produced in some embodiments of the present invention, having a high-methane content, typically at least 80% methane, and with a sufficiently low energy content (BTU/mcf) to be useable in unmodified natural gas engines. This gas may sometimes be referred to as sales gas, or lean methane gas, depending on context, as this is the gas provided or sold to the customer's natural gas gensets for power generation.

B-gas or B-stream refers to the ethane-rich stream produced in some embodiments of the present invention, having a high ethane content, typically 20-60% ethane (with the remainder methane and other hydrocarbon mixtures), and typically used by the ethane-tuned engine for powering the system. The specification and requirements of the B-stream are the most flexible, since the ethane-tuned engine can be tuned to operate on a wide range of ethane levels. This gas may sometimes be referred to as ethane-rich or ethane-enriched gas.

C-liquid or C-stream refers to the natural gas liquids (NGLs) stream produced in some embodiments of the present invention, having a vapor pressure that meets a Y-grade standard for transport in existing pressure vessels or trailers. This stream may sometimes be referred to simply as the natural gas liquids (NGLs) stream.

A distillation column is used to separate a fluid mixture into its constituent parts based on differences in the volatility of components in a boiling liquid mixture. Distillation is a physical separation process and not a chemical reaction. A stripping column is a bottom half of a distillation column. The top portion above the inlet is referred to as a fractionating column and involves an overhead condenser.

Overview of the Mobile Alkane Gas Separator (MAGS)

It makes no financial or environmental sense to flare valuable natural gas liquids that could be sold at great profit. It makes no financial or environmental sense to spend huge amounts of money on diesel fuel for power generation and at the same time flare methane that could be doing the same job. The problem is that liquids-rich raw natural gas cannot be used in generators and cannot be transported by truck. What is therefore needed is a mobile system that can go to the well and separate the raw natural gas into methane that can be used to generate power (or for other purposes) and liquids that can be transported for sale. It is to meet this need that the inventors have developed the Mobile Alkane Gas Separator (MAGS) System.

In the MAGS, raw natural gas is first compressed and then dehydrated. The dry, compressed gas is then refrigerated down to optimally cold temperatures, causing the high molecular weight natural gas components to liquefy. A miniature column system (a stripping column) is then employed to separate the natural gas mixture into three streams. One stream, composed almost entirely of methane with a small amount of ethane, is sent off to power generators to provide electricity for either local use or sale to the grid. Another, composed primarily of ethane and methane (ethane-rich stream), is sent to the MAGS internal power generator to support its own operation. The third, composed of natural gas liquids including propane, butane, pentane, and hexane, as well as higher hydrocarbons, is stored as a liquid so that it can be transported to market for sale.

Therefore, the MAGS solves a long-felt, unsolved need to nearly entirely eliminate flaring by dividing the raw natural gas into three streams, and providing a use for all three streams. This allows oil operators to take flaring to an absolute minimum.

The innovative design of the MAGS according to the principles of the present invention utilizes efficient gas separation using a novel cascade/autocascade refrigeration subsystem. This design allows for compact and cost-effective separation of raw natural gas at the site of an oil field. The natural gas liquids are transported by truck and sold, while the ethane-rich stream is used to power the system itself. Meanwhile, the methane stream is lean, and can be utilized for power generation, high-pressure conversion to CNG which can be used in vehicle applications or transported to remote sites, chemical conversion to liquids such as methanol and higher hydrocarbon liquid fuels, or liquefied into LNG, or for other purposes.

The unique design of the MAGS produces a system that can be sized to fit within the dimensions of a regular semi-trailer. Trailers are limited by U.S. Department of Transportation (DOT) regulations to a size of 13.5 feet by 8.5 feet by 53 feet. The MAGS-200, which can process 200 mcf/day of flare gas, fits onto a trailer sized 13.5 feet by 8.5 feet by 40 feet. A larger unit can be manufactured that processes up to approximately 300-500 mcf/day and still fit within the 53 foot length DOT limit. The trailer weights less than the 60,000 lbs. DOT limit, at approximately 30,000 lbs. in one embodiment of the present invention.

Natural gas at the wellhead is commonly a mixture of methane (C1) with other hydrocarbons, including ethane (C2), propane (C3), butane (C4), pentane (C5), hexane (C6), and higher (C6+). Wellhead natural gas also contains other compounds such as water vapor, hydrogen sulfide, carbon dioxide, oxygen, and nitrogen.

Unmodified natural gas engines are designed to receive preferred methane content of at least 90%, depending on engine manufacturer, corresponding to a methane number of 80. Unmodified natural gas engines can tolerate a methane content of 80%, with higher hydrocarbons (C3+) of no more than 10-15%, depending on engine manufacturer, corresponding to a methane number of 65. Although natural gas engines can be converted to operate with methane contents outside of these ranges, this requires modification of the engine controls and could void the engine warranty. Therefore, engine operators are reluctant to modify their existing natural gas engines. The present invention obviates the need to modify engines since it produces a lean methane stream useable on unmodified engines.

The process described in this patent application has been demonstrated to achieve a methane content in the methane stream of at least 87%, depending on the feedstock. If the raw natural gas stream is 70% methane, the process can achieve over 90% methane. If the raw natural gas stream is 60% methane, then the process can achieve 87% methane. If the raw natural gas stream is lower in methane content, than the process can achieve at least 80% methane content in the methane stream, which is sufficient to meet the minimum requirements of unmodified gas engines.

The processes described in this patent application have been demonstrated to achieve Y-grade NGLs, of approximately 250 psi maximum vapor pressure at 38° C./100° F. As noted in the definitions, Y-grade is not necessarily limited to 250 psi, but refers to a hydrocarbon mixture that has essentially no methane and a low fraction of ethane. Each fractionator has different Y-grade product specifications (pressures and quality parameters of product) and ways of receiving the Y-grade (e.g., pipeline connections or tank car loading capabilities). If a pipeline receiving the NGLs stream can accept 600 psi, then 600 psi vapor pressure would be acceptable in some embodiments of the present invention. The existing propane tanks are rated for 200-250 psia, depending on the transportation company, and most often 225 psia. Therefore, a preferred embodiment of the present invention would achieve the vapor pressure required for transport in propane tanks. The lowest possible NGLs vapor pressure is preferable; however, the lowest achievable vapor pressure in practice is the same vapor pressure as propane, or ˜190 psig, if propane is the predominant NGLs (and possibly lower if it is not present in abundance), since propane has the highest vapor pressure of the NGLs. The lowest vapor pressure possible is desirable because the ethane content brings low value as ethane has a lower price (50.30/gal) than either propane or butane, while taking up shipment weight in the transport container. Therefore, economics are improved when ethane content is reduced.

After the methane and NGL streams, the remaining ethane-rich stream is used internally for powering a specially-tuned ethane engine that powers the entire system. This is in contrast to standard refineries which typically convert the extracted ethane to ethylene for use in chemical synthesis. If too much ethane-rich gas volume is present in the raw natural gas stream, then in one embodiment of the invention, the remaining ethane-rich gas that is not needed for power generation may be flared. Alternatively, the remaining ethane-rich gas may be mixed with the methane stream (A-gas), and the methane stream would still meet engine specifications because there would only be a small amount of the ethane-rich gas left over. If there is not enough ethane-rich gas volume to provide sufficient power to the system, then in one embodiment, some of the methane stream will also be utilized to provide sufficient power for system operations. There are no specific requirements on the ethane-rich gas stream for powering the internal power generator according to one embodiment of the present invention (0-100% ethane is acceptable for the processes described here). The reason this can be achieved is because the onboard engine can be modified and tuned to accept any ethane content. In one preferred embodiment of the present invention, the MAGS power generator is a dual-fuel methane-propane gas engine, which is tuned to run on a variable ethane content, since ethane is intermediate between methane and propane in its energy content.

Potential Economics of the MAGS

One of many illustrative scenarios is presented here to demonstrate the potential profitability of the MAGS system. In this scenario, all of the CH₄ produced is used for electricity production. Other configurations in which the CH₄ is utilized for other purposes is also possible. This economic analysis is illustrative of the invention only and is not meant to limit the scope of the present invention.

In one embodiment, a MAGS-200 field unit processes 200 mcf of raw natural gas per day. Assuming a feed of a NGLs-rich composition typical of many Bakken sites as shown in Table 1, such a MAGS-200 unit can refine such a raw feed into enough methane to produce about 450 kWe of electrical power and about 1700 gallons of natural gas liquids per day. The output produced by such a MAGS-200 unit would vary with the gas input composition.

At current prices, the natural gas liquids produced by a MAGS-200 unit running on such a raw gas feed would have a market value of about $1700 per day, which discounted for transport costs could produce $1000 per day in revenue. If used to replace grid power, the electricity would be worth another $1000 per day, while if used to replace diesel-generated power, $4000 per day in savings could be achieved. The total value delivered by the MAGS-200 unit could thus range from about $730,000 to $1,800,000 per year. Assuming a unit value 4 times that of its annual generated profit, such a unit could therefore have a commercial value between $3 million and $7 million. The revenue from such a MAGS-200 unit would vary with the gas input composition, and would be higher for a gas well with a higher NGLs content.

Preferred System Schematics of the Mobile Alkane Gas Separator (MAGS)

The present invention in its various embodiments provides highly efficient and economic solutions to address important unmet needs to recover and use valuable liquid-rich natural gas. In one embodiment, the Mobile Alkane Gas Separator System is a mobile system that can go to the well and separate the raw natural gas into methane that can be used to generate power and liquids that can be transported for sale.

FIG. 1 shows a schematic diagram 100 of one embodiment of a process employed by the MAGS according to the principles of present invention. Raw natural gas 101 is first compressed, as shown in module 102. The natural gas is then dehydrated to remove any water content, as shown in module 103. The dry, compressed gas is then refrigerated, causing the high molecular weight natural gas components to liquefy, as shown in module 104. A miniature column system (stripping column) is then employed to separate the natural gas mixture into three streams, as shown in module 106. One stream, composed almost entirely of methane with a small amount of ethane, is sent off to external power generators 108 to provide electricity for either local use or sale to the grid 109. Another stream, composed primarily of ethane and methane (ethane-rich stream), is sent to the MAGS internal power generator 105 to support its own operation. The third stream, composed of natural gas liquids 107, including propane, butane, pentane, and hexane, is stored as a liquid so that it can be transported to market for sale.

FIG. 2 shows a block diagram of one embodiment of a cascade/autocascade refrigeration cycle useable in one embodiment of the MAGS according to another embodiment of the present invention. As shown in FIG. 2, the design includes a high stage cascade loop and a low stage autocascade loop arranged and interconnected as shown in FIG. 2. The cascade loop utilizes a single refrigerant, while the autocascade loop utilizes two or more refrigerants. In one preferred embodiment of the present invention, the refrigerants are hydrocarbons. In one preferred embodiment of the present invention, the high stage cascade loop utilizes propylene (also known as propene or methylethylene, C₃H₆), with a boiling point of −47.6° C. In one embodiment of the present invention, the low stage autocascade loop utilizes two or more hydrocarbons. In a most preferred embodiment of the present invention, the low stage autocascade loop utilizes three hydrocarbons comprising: (1) n-butane, having a boiling point of 0.5° C.; (2) propylene, as in the cascade loop having a boiling point of −47.6° C.; and (3) ethylene (also known as ethene, C₂H₄), having a boiling point of −103.8° C. Alternative embodiments of the present invention utilize other hydrocarbon refrigerants, while other alternative embodiments utilize other refrigerants, including non-hydrocarbon refrigerants.

After initial compression and dehydration, as well as initial chilling (not shown in FIG. 2), flare gas enters at flare gas inlet 251 and passes through the high-stage cascade loop (high temperature) at a first flare gas heat exchanger 209 (where flare gas is cooled from approximately 25° C. to approximately 6° C.), and then passes through the low-stage autocascade loop (low temperature) at a second flare gas heat exchanger 204 (where the flare gas is further cooled down to an exit temperature of approximately −57° C.). The refrigerated flare gas then exits the autocascade loop after the second flare gas heat exchanger 204 at flare gas outlet 252 and proceeds to phase separation (not shown in FIG. 2) as described later. Note that various incidental components, such as filters, etc. that have no contribution to the thermodynamic cycle are not shown in FIG. 2 for ease of presentation.

The operation of the cascade/autocascade refrigeration cycles will now be described, starting with the high-stage cascade loop, and secondly the low-stage autocascade loop. Finally, the temperatures and pressures of the flare gas and the refrigerants will be described at each point in the process. The interstage heat exchanger/condenser 210 (also referred to as the auxiliary condenser) connects the high-stage cascade loop with the low-stage autocascade loop. Aside from the interstage heat exchanger 210, the phase separator 205 forms the central core of the system, in which the three hydrocarbon refrigerants of the autocascade loop are phase separated (n-butane and propylene being liquid dominated, and ethylene being vapor dominated in the phase separator 205).

In the high-stage cascade loop, refrigerant (propylene in this example) passes through first compressor 211, where it is compressed and gains heat (heat of compression), after which it passes through an air-blown condenser 212 which reduces some of the heat of compression. The refrigerant then passes through expansion valve 213, where it is cooled substantially via the Joule-Thompson effect. Then, the cold refrigerant from the high-stage cascade loop exchanges heat with the refrigerants in the low-stage autocascade loop via interstage condenser/heat exchanger 210, after which it exchanges heat with the flare gas in the first heat exchanger 209, completing the cycle.

In the low-stage autocascade loop, the multi-component, high-pressure refrigerant is first cooled in the interstage condenser/heat exchanger 210, and then passes to phase separator 205, in which the multi-component refrigerant is separated into its liquid phase (bottom) and gas phase (top). In the phase separator 205, the n-butane and propylene is liquid, having a higher boiling point, while the ethylene is vapor, having a lower boiling point. The liquid portion (n-butane and propylene in this example) passes to a second expansion valve 207, where it expands and cools before entering static mixer 203 (which will be discussed later). Meanwhile, the gaseous portion (ethylene in this example) passes through an autocascade heat exchanger 206, where it exchanges heat with the cold refrigerant from static mixer 203, and condenses into a liquid state. The now liquid ethylene portion then passes through a third expansion valve 208, where it further loses heat via the Joule-Thompson effect. Immediately after expansion valve 208 is the lowest temperature point of the system, as low as −60° C. in some embodiments (possibly −20° C. to −100° C.). Finally, the lowest temperature refrigerant passes through a second flare gas heat exchanger 204, where it cools the flare gas before a flare gas outlet 252. After passing through the second flare gas heat exchanger 204, the ethylene refrigerant mixes with the other two refrigerants (n-butane and prolylene) in static mixer 203. After passing through autocascade heat exchanger 206, the three refrigerants are then re-compressed via a second compressor 202, where the refrigerant gains significant heat (via heat of compression). A portion of the heat is removed via the second air-blown condenser/heat exchanger 201, after which it enters interstage condenser/heat exchanger 210, where it is further cooled, completing the autocascade cycle.

The temperatures and pressures of the flare gas and the refrigerants will now be described at each point in the process. These temperatures and pressures are illustrative of but one embodiment of the present invention, and are not intended to limit the scope of the present refrigeration subsystem. Flare gas enters the flare gas inlet 251 at a pressure of ˜200 psi (˜14 bar) and temperature of ˜25° C. After passing through the first flare gas heat exchanger 209, there is a slight pressure drop of 2-3 psi, and the flare gas temperature is reduced to ˜6° C.

In the high-stage cascade loop, after compressor 211, the propylene refrigerant enters at a temperature of ˜0° C. and a pressure of ˜3.5 bar, and leaves the compressor 211 at a pressure of ˜17 bar and a temperature of ˜100° C., due to the heat of compression. After the first air-blown heat exchanger 212, the temperature of the propylene refrigerant is reduced to ˜37° C., with a slight pressure drop to ˜16.8 bar. At the expansion valve 213, all of the propylene is a liquid, and after the expansion valve 213, the propylene is a mixture comprising about 70% liquid and 30% vapor by mass. After the expansion valve 213, the pressure of the propylene refrigerant has dropped to ˜4 bar, and the temperature is reduced to about −12° C. due to the Joule-Thompson effect. An interesting effect happens when the propylene passes through the interstage condenser 210, where the pressure drops slightly to ˜3.7 bar, and the temperature actually drops slightly to −14° C., due to two countervailing forces—the heat exchange with the autocascade stage and the single-component boiling which occurs in the interstage condenser 210—which implies that the propylene follows its saturation temperature. After the interstage condenser 210, the propylene still has a little bit of liquid left (5% liquid, 95% vapor). When the propylene passes through the first flare gas heat exchanger 209, it is warmed up to its starting temperature of 0° C., and with a slight pressure drop to ˜3.5 bar, hence completing the cycle.

In the low-stage autocascade loop, entering the phase separator 205, the temperature of the refrigerants is about −2° C. at ˜17.2 bar. The liquid phase (n-butane and propylene) drops to the bottom of the phase separator 205, and passes the expansion valve 207, where the temperature is dropped from about −2° C. to about −41° C., as a result of Joule-Thompson cooling as the pressure is dropped from ˜17.2 bar to ˜4.2 bar after the expansion valve 207, which leads to static mixer 203. The vapor phase (ethylene) passes through the autocascade condenser/heat exchanger 206, where the temperature of the ethylene refrigerant drops from about −2° C. to about −27° C., with a slight pressure drop to ˜17 bar as there is a slight pressure drop through the condenser 206. The entire ethylene refrigerant is liquid as it enters the expansion valve 208. At the expansion valve, the ethylene experiences a pressure drop from ˜17 bar to 4.4 bar, resulting in a temperature drop of the ethylene refrigerant from about −27° C. to about −68° C., which is the coldest point of the refrigeration system.

As a result, the flare gas, which comes in at about 6° C. into the second flare gas heat exchanger 204, experiences a temperature drop to about −57° C., as a result of exchanging heat with the optimally cold (−68° C.) ethylene refrigerant, whose temperature rises to −38° C. The resulting flare gas leaves the system from flare gas outlet 252 at about −57° C. and a pressure of ˜12.4 bar.

The liquid phase (n-butane and propylene) refrigerant from expansion valve 207 and ethylene refrigerant from heat exchanger 204 are combined in static mixer 203, where the temperatures are approximately matched, at about −41° C., and −38° C., respectively. After the static mixer 203, the mixed refrigerant has a temperature of approximately −40° C. In the autocascade condenser/heat exchanger 206 the temperature of the refrigerant rises to about −20° C., as it exchanges heat with the ethylene vapor phase from the phase separator 205, with a slight pressure drop to ˜3.8 bar across the autocascade condenser 206. The ethylene refrigerant is liquefied in the autocascade condenser 206. Penultimately, the second compressor 202 compresses the mixed refrigerant from ˜3.8 bar to ˜17.7 bar, resulting in a temperature increase to 90° C. as a result of the heat of compression. The air-blown heat exchanger 201 reduces the temperature to ˜40° C., with negligible pressure drop (still at 17.7 bar). At this point in the process, all of the refrigerants are in gas phase in the air-blown heat exchanger 201. Finally, the refrigerant goes through interstage condenser/heat exchanger 210, where its temperature is dropped from ˜40° C. to about −2° C., with a slight pressure drop to ˜17.2 bar, completing the thermodynamic cycle.

At the interstage condenser 210, there is approximately a 50% liquid, 50% gas composition of the mixed refrigerants. At each expansion valve 213, 207, and 208, there is ˜100% liquid inlet, with an outlet of ˜75% liquid, but with a significant pressure drop (and associated Joule-Thompson cooling). These temperature and pressure estimates are illustrative of but one embodiment of the present invention, and are not to be interpreted as limiting the scope of the present invention. Furthermore, as will be discussed in greater detail later, this hybrid cascade/autocascade refrigeration design is but one embodiment of the present invention, and other refrigeration subsystem designs are also within the scope of the present invention.

Note that the thermodynamic temperatures presented here include a correction term for the semi-hermetic electric motor encasements, which power the compressors and add heat to the refrigerant independent of the heat of compression. Since electric motor are not 100% efficient and generates some heat, and the motor is cooled as refrigerant moves through the compressor, contributing to compressor discharge temperature. Therefore, the thermodynamic temperatures presented above after each compressor include both the heat of compression and the additional motor heat.

FIG. 3 shows a flowchart of a process 300 for separating alkane gases according to one embodiment of the present invention. An embodiment of the present invention is a method for reducing flaring shown in process 300, which starts in step 302. First, a portable alkane gas separator system (MAGS) is brought to an oil field that is flaring associated gas, as shown in step 304. Raw natural gas is received from the flare site, as shown in step 306. The raw associated natural gas stream is compressed utilizing a compressor, as shown in step 308. Water is removed from the raw natural gas stream utilizing a dehydrator, as shown in step 310. Next, a temperature of the raw natural gas stream is lowered utilizing a refrigerator cascade/autocascade loop, as shown in step 312. Next, as shown in step 314, the raw natural gas stream is separated into three product streams consisting essentially of a methane stream of at least 80% methane (and preferably at least 85% methane, and more preferably at least 90% methane), an ethane-rich stream comprising ethane (typically 20-60% ethane), and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C./100° F. (and preferably no more than 225 psia, and even more preferably no more than 200 psia). In step 316, the natural gas liquids stream is of sufficiently low vapor pressure to be transportable in standard propane pressure vessels (Y-grade), where it is collected for transport. In step 318, the methane stream is of sufficient purity and sufficiently lean to be useable in existing natural gas engines without modification, where it is utilized for electricity generation. In step 320, the ethane-rich stream is utilized within the system itself to power system operations. As shown in optional step 322, in some embodiments, the lean methane stream may be utilized to make CNG (via compression), LNG (via liquefaction), or liquid fuels (methanol, butanol, etc.). As shown in optional step 324, the lean methane stream may be utilized with superheated steam to make CO₂ and H₂ via a steam reforming reaction, with the CO₂ useable for enhanced oil recovery (CO₂-EOR), and the H₂ useable to make electricity or for other purposes, such as chemical upgrading of the oil produced. As shown in optional step 326, the electricity may be transmitted to the grid for sale. The process 300 ends in step 328.

FIG. 4 shows a side perspective view 400 of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention. As shown, the system is designed to fit inside the height and width envelope of a standard drop-deck semi-trailer, allowing enhanced mobility to various field sites. The system is mounted on a chassis or skid 404 with one or more wheels 402 for easy transportation to and from the oil field. Raw natural gas is compressed by compression module 406. The natural gas is then dehydrated to remove any water content by dehydration module 408. The dry, compressed gas is then refrigerated, causing the high molecular weight natural gas components to liquefy, by refrigeration module 410. A column separation system (stripping column) 412 is then employed to separate the natural gas mixture into three streams. One stream, composed almost entirely of methane with a small amount of ethane, is sent off to external power generators (not shown) to provide electricity for either local use or sale to the grid. Another stream, composed primarily of ethane (ethane-rich stream), is sent to the MAGS internal power generator 414 to support its own operation. The third stream, composed of natural gas liquids (not shown), is stored as a liquid so that it can be transported to market for sale.

FIG. 5 shows a top view 500 of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention, with similar modules labeled as in FIG. 4.

FIG. 6 shows another side perspective 600 view of one embodiment of the mobile alkane gas separator (MAGS) according to one embodiment of the present invention, with similar modules labeled as in FIG. 4.

In other embodiments of the present invention, the MAGS system is configured and operated so components of the feed natural gas may be separated into more than three product streams or fewer as needed.

Process Parameter Selection

The MAGS process takes an un-useable, stranded, raw natural gas stream, and produces three distinct, useable streams in a field environment. This process is unlike what happens at a refinery, which requires a complete distillation column for each component, resulting in a much larger system and more costly operations. In a refinery, each element would be separated using a de-methanizing column, a de-ethanizing column, de-propanating column, and so forth. This would create a complex, expensive, and cumbersome system that would not be portable and would not be useful in the field. In contrast, the MAGS process is able to achieve a separation as would be achieved at a refinery without using a complex system comprising multiple distillation columns for each component. The inventors were able to achieve this unexpected result by carefully selecting the process operations and process parameters as described in detail in this application.

In summary, the inventors have found that the it is possible to achieve this result by running the refrigeration subsystem to achieve a process temperature preferably in the range of approximately −40° C. to approximately −70° C. (and even more preferably −50° C. to −60° C.), at a pressure preferably in the range of approximately 100 psig to approximately 500 psig (and even more preferably under about 250-300 psig), and that this can be achieved utilizing a novel refrigeration unit. The inventors have discovered that the process parameters described here allow this type of separation to occur at the field scale, something that has not been achieved before in the prior art. By utilizing these process parameters, the inventors were able to optimize the design of the entire system that utilizes this process, allowing the inventors to simplify the system to the point where it can be made portable. The inventors have found that when you compress and cool to the temperature and pressure ranges described here, it is possible to achieve this significant separation result in a portable apparatus.

In one embodiment, the process works by reducing a temperature of the raw, dehydrated natural gas stream to a cold temperature that is an optimally cold temperature range. Unlike a conventional refrigeration cycle, the temperature is colder in order to remove (condense) most of the ethane and higher hydrocarbons. However, it is not so cold (as in LNG processes) in which the methane itself condenses into the liquid stream and would then have to be removed. The MAGS process temperature and pressure parameters were finely calibrated by the inventors in order to achieve this efficient separation process, making it feasible to make the process work on a portable scale. Because of the optimally cold temperature range which condenses most of the ethane, the C2+ components can be removed in a single column. The prior art is either too warm to remove all of the C2+ components in a single separation column, or is too cold, and liquefies some of the methane as well, which is undesirable and also requires an additional separation column.

In addition, because of the optimally cold temperature range, the load on both the compressor and separation column is reduced, reducing system complexity and operating costs. The MAGS process can operate at approximately 200 psi (while the prior art has a higher operating pressure), reducing the compressor load, allowing a smaller and more compact compressor to be utilized. In addition, because of the optimally cold temperature, instead of a complete distillation column, a much simpler stripping column may be utilized. A stripping column is the bottom half of a distillation column, while the top half above the inlet is referred to as a fractionating column and involves an overhead condenser. In a preferred embodiment of the present invention, the system does not utilize a distillation column, nor a fractionating column, but a simpler and less expensive stripping column. (A distillation column is twice as complex as a splitting column.)

The optimally cold temperature range allows the MAGS process to refrigerate most of the ethane and higher hydrocarbons (C2+) out of the gas stream. The MAGS process condenses much more of the ethane than is typical of refrigeration cycles. This allows the liquid stream to be captured and brought to Y-grade standard with just a single column separator; while the lean gas mixture is ready for power generation equipment without any additional processing, greatly simplifying system complex and reducing cost.

In a surprising and unexpected result, the inventors have found that refrigeration is less expensive than compression; and that by reducing the temperature of the raw gas stream to a lower temperature range than in the prior art, a lower compression ratio is possible while still achieving desirable results. The inventors' innovative approach was to realize that by selecting the appropriate temperature range to split the incoming raw natural gas stream into two, one that is predominantly methane (A-stream), and the balance (mixture of B-stream and C-stream) which can be separated in only a single column (into B-stream and C-stream).

This process greatly saves on the capital costs, as it eliminates the cost and complexity of an entire column (one column is sufficient instead of the typical two); reduces compression requirement, typical prior art applications would require 3-5 stages of compression, (whereas the MAGS process only requires 2 stages of compression, saving on compression equipment costs); reduces the operating pressure, as typical prior art systems utilize up to 1,000 psi (whereas the MAGS process only requires 100-300 psi), resulting in a simpler and cheaper downstream system that doesn't need to operate at such high pressures.

The temperature and pressure ranges for the MAGS process were selected as follows. A most preferred temperature range for the process is −50° C. to −60° C., but temperature ranges from −20° C. to −100° C. are feasible, and more preferably between −40° C. and −70° C., and possibly between −20° C. to −40° C. There would be little benefit to going below −100° C. and a significant cost to go below that temperature, as methane starts to liquefy below −100° C., and it is desirable to avoid liquefying any significant amount of methane. Accordingly, based on the above temperature ranges, a most preferred pressure range is 175-200 psi, and preferably 150-300 psi. However, pressure ranges of 100-500 psi are possible in various embodiments of the MAGS process. Higher pressures are also possible in some embodiments. Prior art processes must utilize higher pressures because they have not selected optimal process temperatures.

In addition, many prior art processes relied on higher compression and expansion cooling to refrigerate, and did not perform active refrigeration as in the MAGS process. Unexpectedly, the inventors found active refrigeration to offer many advantages, as described here. On a high level, the MAGS processes utilizes compression, refrigeration, and separation. If we assume that each of the three sub-processes have been optimized and has equal cost and complexity, what this implies if that optimizing one of the three sub-processes it is possible to save on the capital and operating costs of the other sub-processes. The inventors had the insight that by improving the middle of the three processes—namely, refrigeration—it was possible to significantly save on the cost and complexity of the two processes before and after it—namely, compression and separation. After the raw, dehydrated natural gas mixture is compressed and refrigerated to the optimized pressure/temperature ranges described above, there is a simple liquid/vapor separation. The gas stream (A-stream) that leaves the liquid/vapor separator is ready to be used in industrial equipment without further processing. Meanwhile, the remaining liquid stream can be cleaned up in a single column into a NGLs stream (C-stream), and a remaining waste stream (B-stream), which can be used by the system itself for power generation. That is, after compression and refrigeration, the stream components that remain after phase separation and that go to the splitting column are primed for separation in as few as a single column.

In summary, the prior art is able to achieve two-stream separation. The prior art either generates lean methane and flares the NGLs, or generates NGLs and flares the methane and ethane. By carefully selecting the process parameters and utilizing a novel system design, the inventors have been able to achieve something on the portable scale that before was thought only possible at a refinery scale.

Experimental Results

Several random samples of Bakken region wellhead gas quality data is presented in Table 1 (Source: Wocken, C. A.; Stevens, B. G.; Almlie, J. C.; Schlasner, S. M., End-Use Technology Study—An Assessment of Alternative Uses for Associated Gas, National Energy Technology Laboratory, Pittsburgh, Pa., April 2013). This high NGLs content typically corresponds with high Wobbe index (higher energy content of 1300-2000 BTU/cf) when compared to residential pipeline gas (1000 BTU/cf).

Based on the sample Bakken gas data from Table 1 as well as a sample assay provided by a major oil company in the Bakken region, a simulated Bakken gas consisting of 60% methane, 18% ethane, 14% propane, and 8% butane was generated in the lab for testing. Table 2 shows the latest experimental results, using a 12 mcf/day subscale demonstration unit (“MAGS-0”), where Table 2A shows a predicted results based on computer simulations in HYSYS, and Table 2B shows the actual experimental results on the prototype unit. All flow rates are normalized to 100%. The actual flow rate was 240 L/min of simulated raw feed gas. As can be seen from Tables 2A and 2B, there is excellent agreement between the theoretical predictions and the experimental results of running the prototype unit. A lean methane gas stream of 88% methane was generated (A-gas), which can be utilized in unmodified natural gas engines. The natural gas liquids stream (C-liquid) was also produced having a low enough vapor pressure sufficient for transport in propane tanks. Note that 81% of the propane and 99% of the butane liquids were captured in the NGLs stream. Finally, the ethane-enriched gas stream (B-gas) was produced having a 44% ethane, 50% methane, and some residual propane and butane which can be utilized in the internal MAGS power generation subsystem as described below.

TABLE 1 Select Associated Gas Quality Data from Wellheads in the Bakken Formation Wellhead Sample: Sample 1 Sample 2 Sample 3 Methane, mol % 70.23 48.07 73.93 Ethane, mol % 13.94 18.78 13.25 Propane, mol % 6.7 14.87 5.55 Butane + (C4+), mol % 5.5 16.38 4.32 CO₂ + N₂, mol % 3.44 1.72 2.87 H₂S 0.19 0.18 0.08 Wobbe Index, Btu/scf 1470 1712 1454 Methane No. (MN) 53.2 43.5 56.1

TABLE 2A Theoretical Predicted Performance using HYSYS, MAGS-0 MAGS-0 Theoretical Predicted Performance Gas Gas NGL Feed Prod- Prod- Prod- Gas uct A uct B uct C % of NGL Product Products (%) (%) (%) (%) Captured Methane 60.0 87.5 42.9 0.9 Propane Butane Ethane 18.0 10.6 43.2 22.2 81.8 97.2 Propane 14.0 1.7 11.6 45.8 Pressure: 180 psia Butane 8.0 0.2 2.3 31.1 Refrigerate Temp.: −55° C. Total 100.00 62.30 12.70 25.00 Ambient Temp.: 25° C. flow rate

TABLE 2B Experimental Results, MAGS-0 MAGS-0 Experimental Results Gas Gas NGL Feed Prod- Prod- Prod- Gas uct A uct B uct C % of NGL Product Products (%) (%) (%) (%) Captured Methane 60.1 88.7 49.6 0.4 Propane Butane Ethane 18.0 9.9 43.7 16.7 80.7 99.5 Propane 13.6 1.4 6.0 47.3 Pressure: 180 psia Butane 8.3 0.0 0.7 35.6 Refrigerate Temp.: −55° C. Total 100.00 57.70 19.10 23.20 Ambient Temp.: 25° C. flow rate

Detailed Schematics of the Mobile Alkane Gas Separator (MAGS)

A detailed process flow diagram for the separation of the three products is now described according to one embodiment of the MAGS shown in FIGS. 7A and 7B. The process flow diagram depicts the flow of raw, wet natural gas through the various components of the MAGS process. Gas is obtained from the wellhead after oil, water, and sand have been removed by a heater-treater separation unit (not shown).

The MAGS is designed to receive wellhead natural gas 701 from a remote field location. The MAGS system processes the raw natural gas into three component streams: lean methane-rich dry sales gas (A-gas) intended for client-side power generation 717, high-energy content ethane-rich dry gas (B-gas) to generate power required by the MAGS system 719, and a market-ready Y-grade natural gas liquids (NGLs) product 711 (C-liquid).

The raw natural gas processing begins with compression at 702. The wellhead natural gas 701 received from heater-treater units (not shown) enters the system at moderate pressure and is regulated down to a consistent pressure (˜20 psig) at 702 a. The first stage of compression 702 b achieves a moderate pressure rise and is accompanied by an air-cooled inter-cooler 702 c and liquid knockout at condenser/dehydrator 702 d. The second stage of compression 702 e achieves the final operating pressure required by downstream processing. An after-cooler 702 f and liquid knockout at condenser/dehydrator 702 g ensure no liquid condensate enters the downstream process. In this embodiment, two stages of compression are used to bring the raw natural gas to a sufficiently high pressure to enable separation of the constituent components. Each individual stage of compression is accompanied by an air-cooled heat exchanger to reduce the temperature of the stream and to condense any humidity present (first stage of dehydration). After cooling, a liquid condensate separator removes the liquid water from the natural gas. In short, compression occurs in two stages, where the gas stream is cooled to ambient temperatures after each compression stage, during which water and heavier hydrocarbons (C6+) will drop out (as liquids), before the final dehydration step (desiccant beds). Most of the water is condensed as a result of the inter-cooler 702 c and after-cooler 702 f in the first two condensers/dehydrators 702 d and 702 g, so that most of the water will be removed before the final dehydration step (desiccant beds). Also, most of the hexane, octane, etc. are knocked out during this stage.

The compressed natural gas enters a regenerative pressure-swing adsorption desiccant system 703 for the final dehydration that removes any water remaining after the first two stages of dehydration in the condensers 702 d and 702 g. The resulting natural gas has an aqueous dew point below −73° C. (lower than the coldest temperature in this embodiment) which reduces the potential for hydrate formation in the remaining portion of the system. The maximum dew point can be adjusted based on the system's operating temperature at its coldest point (after the autocascade). One can see from FIG. 7A that there are two desiccant beds 703 in parallel, allowing one to be regenerated by dry, hot gas from other parts of the process while the other bed is valved into the main process line and is actively drying the raw natural gas.

Once the gas is sufficiently dehumidified and at high pressure, the natural gas is now ready to begin the chilling process to condense out natural gas liquids at the elevated operating pressure. The optimally cold temperature is achieved by a hybrid cascade-autocascade refrigeration subsystem, described previously in greater detail in relation to FIG. 2. The elements shown in FIG. 2 are also illustrated in FIG. 7 but with a different orientation (for example, the phase separator 205 from FIG. 2 corresponds with the phase separator 714 in FIG. 7). First, heat exchanger 704 counter-flows the cold, separated A-gas stream with the raw natural gas stream, as described below. Secondly, the raw natural gas stream is cooled again by heat exchanger 705 located on the high-stage cascade refrigeration loop 712. Thirdly, the raw natural gas stream is brought to the design temperature required for liquefaction of natural gas liquids at the system operating pressure by heat exchanger 706 which is cooled by the low-stage autocascade refrigeration loop.

The first refrigeration stage, the high-stage refrigeration loop 712, incurs load from the low-stage autocascade refrigeration loop as well as the dehumidified raw natural gas at the heat exchanger 705 (corresponding to the flare gas heat exchanger 209 in FIG. 2). This is the initial cooling step for the process gas. The natural gas is chilled to its lowest process temperature at the second refrigeration stage, the low-stage autocascade loop, at heat exchanger 706 (corresponding to heat exchanger 204 in FIG. 2).

As described previously, the low-stage autocascade refrigeration loop achieves its optimally low temperature by employing a mixture of refrigerants. The refrigerants are compressed by compressor 713 and partially evaporated by the low-stage refrigeration cycle 712. The two-phase mixture enters a phase separator 714 where the vapor exiting is cooled by condensation of the low-vapor pressure refrigerant components as well as heat exchanged with the recombined refrigerant stream 715. This scheme results in the ability to achieve an optimally low temperature at the heat exchanger 706 (corresponding to second flare gas heat exchanger 204 of FIG. 2), permitting a process gas temperature as low as −60° C., or even −100° C. in some embodiments.

An elevated pressure and reduced temperature causes a portion of the higher-hydrocarbon components of the natural gas stream to condense out. The vapor-liquid separator 707 separates the two-phase natural gas stream into a lean methane-dominant sales gas stream (A-stream) and a natural gas liquids stream comprising of C2+ hydrocarbons (what later is separated into the B-stream and C-stream). This vessel prevents entrainment of liquid droplets with the sales gas (A-stream) leaving the top, thus maximizing capture of the higher-vapor pressure components in the NGLs. This vessel also doubles as a feed drum for the de-ethanizer stripping column 708.

The stripping column 708 is required to lower the vapor pressure of the Y-grade natural gas liquids product. The stripping column 708 is a vessel that separates compounds based on their difference in vapor pressures. The column 708 is packed with mass transfer material that facilitates the exchange of components between the liquid and vapor phases. There is a temperature and compositional gradient along the vertical axis of the column. The top of the column is dominated by a low-temperature, low-boiling vapor, while the bottom of the column is at a higher temperature and consists predominantly of high-boiling liquids. The stripping column permits light components to leave the top through a pressure-control valve, while the liquids exit the bottom of the column to a reboiler 709 (such as a kettle reboiler).

The reboiler 709 provides the heating duty necessary to drive the lighter components from the top of the column as a vapor stream. Heat is added to vaporize lower-boiling components to disengage from the liquid product stream (bottom) and reenter the column. Buoyancy forces and packing material facilitate countercurrent heat and mass transfer inside the column between the liquid and vapor phases. This achieves the design goal of obtaining a final natural gas liquids product that meets the Y-grade specification of having a moderate vapor pressure at ambient temperatures. The reboiler 709 also serves as one of the primary degrees of freedom in the control scheme, allowing the system to handle inlet feed gases of varying composition.

The liquid level control pump 710 ensures a liquid level is maintained inside the reboiler 709. The pump 710 also provides a pressure head to pump the NGLs product (C-stream) into liquids storage tank 711. The NGLs (C-liquid stream) can be transferred from storage to truck trailers for transport to market.

The higher energy content dry natural gas (ethane-rich B-gas stream) that leaves the top of the stripping column 708 is at moderate temperature so no further heat exchange is required. It is simply transported to the fuel pressure regulator 718 and dropped to a low pressure for use in the onboard genset 719. The gas is combusted in an internal engine 719 a and the mechanical shaft power is transmitted to a 480 VAC internal generator 719 b operating at 60 Hz. This power is then available to drive the compressors and electrical components onboard the MAGS system, as shown in dashed lines.

The dry sales gas (lean methane A-gas stream) that leaves the top of the vapor-liquid separator 707 is very cold and at the system pressure, so it is efficient to recapture some of this energy before the gas leaves the system. The methane stream drops some pressure through the level control valve at the top of phase separator 707 before chilling the dehumidified process gas leaving 703 at heat exchanger 704. The methane stream is then heated to a significantly elevated temperature by absorbing the energy from the onboard genset exhaust stream at heat exchanger 716. This hot gas is then used to regenerate the offline desiccant bed at 703 by removing moisture from the adsorbent material. This gas is then cooled again by exchanging heat with liquid water from the condensers 702 d and 702 g at heat exchanger 702 h. Finally, the gas is cooled again with an air-cooled heat exchanger and its pressure is dropped to an operating pressure suitable for use in the client's existing onsite generators 717.

The lean methane product (A-gas stream) produced by the MAGS process is generator-quality methane of a sufficiently low BTU content to be useable in standard, readily available, unmodified industrial gensets. The methane-rich fuel can be burned in an industrial genset 717 to replace the cost of on-site power generation required to run other systems at the wellhead, or for other purposes such as vehicle CNG, as described in the present application. The lean methane stream comes out at a positive pressure of up to ˜4-6 bar, or ˜50-100 psi. This moderate pressure methane stream can be either reduced down to ambient for use in industrial generators or compressed up to high pressures (˜2,000-3,000 psi) for CNG use.

In some alternative embodiments of the present invention, multiple phase separators A, B, C, etc. (not shown in FIG. 7B) can be utilized in addition to the single phase separator A (707) shown in FIG. 7B. For example, the liquid stream leaving a first phase separator A flows through a restriction which drops both the temperature and the pressure as the liquid undergoes adiabatic expansion and Joule-Thompson cooling. The output enters a second phase separator B, where another pair of liquid and vapor phases are separated. As before, the liquid phase is enriched in the heavier, higher-boiling natural gas components, while the vapor is composed primarily of lighter, lower-boiling natural gas components. The exact number and specific conditions of these phase separators is a variable of this process. An additional restriction and a subsequent phase separator C are possible. The final phase separator C would also serve as a liquid feed accumulator for steady operation of the stripping column 708. If multiple phase separators are utilized, then their output is combined with the output of the stripping column 708 to create a single internal ethane-rich stream.

In some alternative embodiments of the present invention, the final natural gas liquids product is delivered to one of two parallel liquid receiver tanks (not shown in FIG. 7B). One is filling with the low pressure, low temperature product, while the other is being heated to raise its pressure. This allows the offline tank 711 to develop sufficient pressure to be delivered to bulk storage without utilization of the pump 710, saving some energy by eliminating the pumping mechanism at the cost of having two liquid receiver tanks.

Refrigeration Subsystem Embodiments

Since temperatures lower than −40° C. are preferred in this invention, a cascade or autocascade refrigeration systems is preferable to be used. In some embodiments of the present invention, a novel cascade/autocascade hybrid refrigeration system is utilized, as shown and described in relation to FIG. 2. Some of the features and advantages of such a cascade/autocascade hybrid refrigeration system is now discussed, along with alternative embodiments. In a refrigeration system, the maximum difference between the warm and cold temperature of a refrigeration cycle is limited by properties of the refrigerant and/or losses associated with the transport of the refrigerant. For larger temperature differences, one has to arrange several refrigeration cycles “above” each other, each cycle spanning a certain temperature difference. A cascade refrigeration system is made up of separate but connected refrigeration stages, each of which have a primary refrigerant wherein the refrigerants work in concert to reach the desired temperature. The principal is to condense refrigerants that are capable of achieving ultra-low temperatures that would otherwise not be able to condense at room temperature. The reason that two refrigeration stages are used is that a single stage cannot economically achieve the high compression ratios necessary to obtain the proper evaporating and condensing temperatures.

The cascade refrigeration system comprises two separate circuits, each using a refrigerant appropriate for its temperature range. The two circuits are thermally connected by the cascade condenser, which is the condenser of the low-temperature circuit and the evaporator of the high-temperature circuit. Refrigerants that may be selected for the high-temperature circuit include R-22, ammonia, R-507, R-404a, and so forth. For the low-temperature circuit, a high-pressure refrigerant with a high vapor density (even at low temperatures) should be selected (such as ethylene).

The condenser of the first stage A, called the “first” or “high” stage, is fan cooled by the ambient air. In some embodiments, a water supply may be used. The evaporator of stage A is used to cool the condenser of stage B, called the “second” or “low” stage. The unit that makes up the evaporator of stage A and the condenser of stage B is referred to as the “inter-stage” or “cascade condenser.” Cascade systems use two different refrigerants, one in each stage. The two-stage cascade system uses these two refrigeration systems connected in series to be able to achieve temperatures of around −85° C.

In an auto-cascade refrigeration system, in contrast, a single compressor system that can achieve temperatures as low as −100° C. is utilized. An autocascade refrigeration system is a complete, self-contained refrigeration system in which multiple stages of cascade cooling effect occur simultaneously by means of vapor-liquid separation and adiabatic expansion of various refrigerants. Physical and thermodynamic features, along with a series of counterflow heat exchangers and an appropriate mixture of refrigerants, make it possible for the system to reach low temperature. The autocascade is a cooling/freezing system using one compressor and two or more different refrigerants and heat exchangers to reach a lower temperature, wherein the first refrigerant cools the next and so on. The components of an autocascade refrigeration system include a vapor compressor, an external air- or water-cooled condenser, a mixture of refrigerants with descending boiling points, and a series of insulated heat exchangers.

The autocascade refrigeration system uses mixed refrigerants along with internal heat transfer and phase separation to achieve optimally cold temperatures through a single compressor. The most basic autocascade cycle uses only a single phase-separator and one additional heat-exchanger (compared to a standard refrigeration cycle) to mimic the behavior of a conventional two-stage cascade refrigeration system.

The refrigerant in the autocascade cycle is compressed as a gas and then sent through a condenser where heat is removed to liquefy the refrigerant. Because an autocascade uses mixed refrigerants of differing vapor pressures, the condensation of the gas is only partial. The refrigerant with the higher vapor pressure remains predominately gaseous whereas the refrigerant with a lower vapor pressure is liquefied. This two phase flow is then sent to a vessel where the gas and liquid phases are separated. The liquid stream is dropped in pressure to provide a cooling effect which is used—in a heat-exchanger—to further chill and condense the gas stream. The gas stream (now liquefied) is then dropped in pressure to provide the final useful cooling duty desired. In this way, an autocascade cycle essentially replaces two stages of conventional cascade refrigeration.

The above description is of the simplest autocascade cycle possible. Significantly more complex cycles are possible for use with the present invention. In some embodiments of the present invention, additional “staged” phase-separation steps with their corresponding internal heat transfer cooling afterwards can be used to reach even colder temperatures.

As a result of its multiple refrigerants and unique design, the autocascade design of the present invention can attain deeper temperatures in a single stage than possible in a conventional cascade refrigeration cycle, so much so that a single stage autocascade refrigerator can replace a two-stage (or more) cascade refrigerator. While such an autocascade unit would not be as energy efficient as a two-stage conventional cascade system, it would be simpler and cheaper to build and operate. Because the MAGS typically operates in an energy-rich environment, the trade of reduced capital and operating costs at the expense of increased energy costs offered by the autocascade is potentially highly attractive and may be considered a preferred embodiment.

The temperatures reached by the autocascade may be altered by altering a composition of the mix of refrigerants. Depending on the composition of gas at the wellhead, the MAGS system may thus be tuned to reach appropriate temperatures for effective operation with the gas composition at hand.

In a preferred embodiment, the MAGS refrigeration system is an innovative cascade/autocascade hybrid, as discussed previously in relation to FIG. 2. This allows for a more thermodynamic efficient design while keeping the system compact and portable. While the thermodynamic efficiency is not necessarily optimal, because the MAGS operates in an energy-rich environment, it is sensible to trade some thermodynamic efficiency for a more mechanically compact design.

In the cascade/autocascade hybrid, two stages of refrigeration are employed, with the first stage cooling the gas to just above 0° C., and the second stage being used to chill it to much colder temperatures, typically ranging from −40° C. to −60° C., depending upon the particulars of the design. In the preferred configuration, the first (warmer) stage utilizes two evaporators, while the second (colder) stage is of an autocascade design. Both stages utilize air-cooled heat-exchangers to eliminate the need for water which may not be available at all operating sites. However, if water is available, the refrigeration cycle can be modified to utilize this resource for enhanced refrigeration performance.

In one alternative embodiment, the first stage is similar to the warmer stage of a standard cascade refrigeration cycle (but with two evaporators) and the second, colder stage, is also a cascade design.

For improved thermodynamic efficiency, both refrigeration loops may be used, in series, to chill and condense natural gas liquids from the natural gas stream. This requires two evaporators on the warmer refrigeration loop (the cascade evaporator/condenser heat-exchanger and a second heat-exchanger whose duty chills the natural gas stream as well as provides some superheat to the vapor returning to the refrigeration compressor suction inlet) as described in relation to FIG. 2. Various alternative designs can range from a simple, single stage refrigeration cycle to three or more stages of cooling. Further any stage of a single or multi-stage configuration could be of a cascade design, or alternatively, an autocascade refrigeration stage, or alternatively of the hybrid design according to the present invention.

Dehydration Subsystem

Dehydration is necessary to remove entrained water moisture from the raw natural gas stream before refrigeration to avoid ice formation, which would damage or destroy equipment. In a preferred embodiment, dehydration occurs first in the after-cooler condensers, where most of the water is removed after compression and air-cooling. Final dehydration occurs in desiccant dryer beds. Examples of desiccants used in the desiccant beds include silica, alumina, silica alumina, calcium oxide, molecular sieves (such as zeolites), activated charcoal/carbon, and other like materials. Other examples of desiccants which are useable with the present invention include montmorillonite clay, calcium chloride, and calcium sulfate.

Sometimes the water is not removed directly, but additives such as methanol or ethyl-glycol are sprayed into the natural gas stream to prevent ice from nucleating. However, when these NGLs end up at refinery, the refinery has to remove the additives and water. By not introducing such additives into the NGLs stream, a preferred embodiment of the process produces a more valuable NGLs stream by removing water by dehydration without the use of additives.

Dehydration is typically carried out in the prior art using a 4-step process (less effective operation, needing larger beds, and lost time). The beds of the present invention in its preferred embodiment are smaller and more efficient because the inventors have developed a novel 2.5 step dehydration process. The prior art 4-step process includes: (1) actively dehydration, (2) depressurizing, (3) regenerating, and (4) re-pressurizing. If 4 steps are used, then 4 beds are needed. But the inventors have discovered a novel dehydration process using only 2 beds by optimizing the cycle into only 2.5 steps. The 2.5 steps of the dehydration process include: (1) actively dehydrating, (2) regenerating, and a ½-step re-pressurizing cycle, which is quick, but enough to operate with only 2 beds. A small 2-way valve is used between the two beds at the bottom to equalize the pressure (valve 1311 in FIG. 13). Finally, 4-way valves simplify switching between the two beds in the 2.5 step process (valves 1302 and 1304 in FIG. 13).

Adsorbent molecular sieves can achieve moisture contents as low as 0.1 ppm, thereby mitigating the risk of damaging cryogenic components such as pipes, heat exchangers, and expansion devices by freezing water inside them. The molecular sieve material is typically distributed inside round vessels in a packed bed configuration. Like other desiccants, molecular sieves have limited adsorption capacity and must be replaced or regenerated at given service intervals. For continuous dehydration service, a multi-bed system must be utilized where one bed is in service while the other is being replaced or regenerated, and the beds can be seamlessly switched in and out of service.

In general, alternating two-bed systems are used where bed “A” is in service and the process stream is dehydrated. At the same time, a dry regeneration gas is flowed through bed “B” in a counter-current direction to remove moisture from the surface of the adsorbent material. Once the regeneration is complete, a set of valves are actuated such that the process gas is directed into bed “B” and the regeneration gas is flowed through bed “A” counter-current to the process flow. This cycle can be repeated indefinitely until the adsorbent exceeds its useful life, usually years.

Typically, the adsorbent beds are sized so that cycle times are on the order of hours. The packed bed diameter is tuned to provide an acceptable superficial velocity, and the height is adjusted to achieve the required holding capacity. The diameter is limited by pressure containment, and the bed height is limited by overall pressure drop and/or crush strength of the adsorbent material. Optimal sizing can be iteratively obtained by balancing the time required for regeneration with the time available to adsorb water before the holding capacity is reached.

Most regenerative dehydration units employ a temperature-swing process, where the regeneration gas is externally heated. The regeneration gas must carry enough energy to bring the adsorbent material to an elevated temperature, as well as to provide the heat of desorption of the water mass. Additional heat is required to overcome the thermal losses through the piping, vessel wall, and effluent gas. After the removal of the water at the regeneration temperature, the external heater is taken off line and the regeneration gas cools the bed back to the process temperature.

A standard valving arrangement requires four on-off valves per bed, to allow the process stream and regeneration gas to flow through one bed at a time in a counter-current fashion. During the switchover, all eight valves are actuated simultaneously to swap beds. If a pressure difference exists between the process stream and the regeneration gas, a pressure equalization valve between the two beds is required. Pressure equalization must be done gradually to avoid adsorbent attrition, adding to cycle time. If the process gas is not compatible with the regeneration gas, vent valves and inert purge valves may be required to expel the unwanted gas and condition the beds prior to pressure equalization and/or switchover.

In a preferred embodiment of the present invention, two 4A molecular sieve adsorbent beds are used. The process stream is a heavy hydrocarbon gas mixture containing up to 2% water by volume, at a maximum volumetric flow rate of 230 mcf/day. The design inlet conditions are 46° C. and 180 psig; however the unit can operate satisfactorily at off-design conditions. The vessel is sized to provide a maximum superficial velocity of 35 ft/min at flow rates up to 230 mcf/day. The packed height is 2.74 m, resulting in a pressure drop of 3.3 psi and a cycle time of 6 hour at the maximum flow rate. Longer cycles are possible at lower flow rates.

In the preferred configuration, the regeneration gas is the vapor stream off the system phase separator, which is a light hydrocarbon mix at 100 psig and 38° C. The regeneration gas exchanges heat with the system power generator engine exhaust, and is heated to 400° C. before entering the saturated adsorbent bed. The regeneration gas gives up thermal energy to the adsorbent at a rate of 12 kW, leaving the bed at 270° C. After 4.5 hour, the regeneration gas is directed away from engine exhaust heat exchanger, and the bed is cooled from 260° C. to 38° C. by the regeneration gas in 1 hour. There is 0.5 hour of standby time to execute the switchover between beds.

The process gas flows downward through the bed and the regeneration gas flows upward, lifting the adsorbed water from the bed. The system utilizes a simplified valve arrangement based on two 4-way cross-port valves (valves 1302 and 1304 of FIG. 13). The process gas flows through one circuit of the valve while the regeneration gas flows through the other circuit in a counter-current direction. With this configuration, the function of 8 simple on-off valves can be replicated by two 4-way valves placed at the entrance and exit of each bed. At the time of switchover, both valves turn simultaneously. Because both the process gas and the regeneration gas are hydrocarbons, no vent and purge step is necessary.

FIG. 13 shows a schematic diagram 1300 of one preferred embodiment of the dehydration desiccant bed subsystem according to one embodiment of the present invention, in which process gas 1301 from the upstream part of the process is dehydrated to produce dehydrated process gas 1324, which is then sent to the refrigeration subsystem downstream. Meanwhile, A-gas 1310 generated upstream in the process enters via valve 1321 and is utilized to regenerate the dehydration beds before being sent to a customer's genset as sales gas 1320 via valve 1319. As shown in FIG. 13, two dehydration beds 1306 and 1308 are used, along with two 4-way valves 1302 and 1304 to allow simplified switching between the bed in service (1308) and the bed being regenerated (1306). The 4-way valves 1302 and 1304 direct the process gas 1301 into the bed that is in service while simultaneously directing the A-gas 1310 to the bed being regenerated. A 2-way valve 1311 is used between the two beds at the bottom to equalize the pressure. Also as shown in FIG. 13, exhaust gas 1316 from the MAGS engine (not shown) exchanges heat with the A-gas 1310 in heat exchanger 1312 before being vented to the atmosphere 1314. Also shown in FIG. 13 are multiple temperature and pressure sensors 1305, 1307, 1309, 1311, 1313, 1315, 1317, 1318, etc. which are utilized for process control.

In summary, a preferred embodiment of the final dehydration subsystem is an alternating two-bed system, able to dry up to 230 mcf/day of gas in 6 hour cycles. The system has several unique features that save on capital expense and conserve energy. Use of 4-way valves instead of on-off valves simplifies piping and controls while saving space and expense. Utilization of engine exhaust heat to drive the regeneration eliminates the need for expensive combustion or electric heat. Additionally, use of pressurized hydrocarbon vapor as the regeneration gas eliminates a blower, while making the switchover process faster and more seamless. Finally, the system is designed to fit inside the height and width envelope of a standard drop-deck semi-trailer, allowing enhanced mobility to various field sites.

Alternative embodiments of the MAGS dehydration could employ systems involving more than two beds, and/or use other methods of moisture capture, including alternative desiccants, or water capture using coolers or freezers.

Power Generation Subsystem

The MAGS internal power generation unit can run on anything from 0 to 100% ethane. However, it is preferable that most of the ethane is in the B-gas stream (the one that powers the system) so it is not in the NGLs product (C-liquid) or in the A-gas (methane-dominated) stream. The result is that much of the ethane is utilized to power the MAGS system itself so that it does not increase the vapor pressure of the NGLs product nor interfere with engine performance of the lean methane stream. As can be seen from Table 2, about a third (˜32%) of the ethane from the raw gas stream ends up in the A-gas stream, about a sixth (˜22%) ends up in the C-liquid stream, and about a half (˜46%) was used for powering the MAGS itself in the B-gas stream. If a particular raw gas stream has too much ethane content, then in some embodiments, some of the B-gas stream that is not consumed to power the MAGS itself can be mixed with the A-gas stream, while still keeping the A-gas stream above the minimum methane %. Alternatively, in some embodiments, any remaining B-gas stream that is not consumed to power the MAGS can be flared.

The genset used to power the MAGS system is adapted to any fuel mixture (methane/ethane/propane) by using the emissions system to verify stoichiometric operation. The genset varies fuel pressure to achieve the proper air to fuel mixture, wherein the unit varies the time the injectors remain open to meet combustion requirements. Most gensets would not be compression ratio limited, as long as the ignition timing can be varied. There could be a 5-15% derating of the genset's power output, depending upon its propane and natural gas rating. Since ethane resides between methane and propane, safe operation is expected. The power output limiting factor when changing fuels usually comes down to the engine's compression ratio and fuel injection control scheme. In one embodiment, the MAGS system prototype is using a 150 kW rated genset; however, only 65% power capacity (about 100 kW) is required, so derating of the system is not a concern.

For the MAGS-200 unit, a single, 400 kW natural gas genset would be used. It could also provide power for on-site use as well as the system components. For initial start-up, the unit could run off the wellhead gas at a deration until the MAGS system is fully operational. This would also provide a replacement for any rented diesel equipment on-site.

Controls Subsystem

Any chemical process operating real-time in the real world requires a degree of control to maintain process conditions within acceptable operational constraints determined by economic, practical, and safety requirements. The MAGS system employs a novel control scheme in which both stable operation and agility of response to changes in inlet conditions are achieved with minimal complexity and cost.

The first places in the MAGS process allowing for a degree of capacity control are the two stages of compression which increase the pressure of the raw natural gas stream. In the preferred configuration, durations of low flow are augmented with recycled vapor from various parts of the downstream process to maintain constant throughput of the compression equipment. In an alternate configuration, variable frequency drives (VFD) can be installed on the compressor motor so the system can track a variable inlet flow from the wellhead.

Air-cooled heat exchangers rely primarily on convective heat transfer to cool or heat a process stream. Accordingly, the fan speed of each of the five air-cooled heat exchangers in the MAGS process provides a significant degree of control to maintain process streams at desired temperatures.

The preferable configuration of the control scheme relies on static control elements, including both sharp-edged orifice plates and capillary tubes, to maintain the appropriate temperatures and pressures required for separating raw natural gas into three useful product streams. These static control elements allow for a range of operational capacity since only the liquid levels in the phase separators will vary. However, with orifices on the vapor streams and capillary tubes on the liquid streams, physical flow constraints prevent the vessels from emptying or flooding. This can be confirmed with analog level monitoring.

Alternatively, for systems requiring an even higher degree of turndown, such as those installed in a wellhead with great variability in flow, it is possible to install active control elements. These include pneumatically actuated globe valves for vapor streams and electromechanically actuated expansion valves for liquid stream. Both globe valves and expansion valves share the robust level control characteristics of the static control elements discussed above.

Another innovative control feature of the MAGS process was introduced to handle the highly efficient use of thermal integration in the process design. The high-methane vapor stream (A-gas) leaving the first phase separator is heated by the engine exhaust from the onboard generator. In order to deal with variable flows through this heat exchanger, the hot engine exhaust can be diverted with a three-way butterfly valve to avoid overheating the high-methane content stream and damaging valving associated with the desiccant unit.

Local installation of the controls system exhibits redundancy to maintain proper communication between delivery of all three product streams: high-methane fuel, high-ethane fuel, and natural gas liquids. In one embodiment, it is possible to use individual, autonomous pieces of microprocessing hardware to maintain optimal operation of the MAGS systems. The individual control units handle high-level operational control goals, low-level PID loops, communication with both local and remote human operators, and communication with both local and remote MAGS systems, as well as ancillary systems including bulk storage, onboard power generation, and external client-facing power generation.

Communication between hardware is made redundant with backup power supplies, two redundant wired Ethernet networks, and a failsafe wireless network. Local operational highlights are broadcast externally from the field installation via Ethernet radios and satellite communication.

Theoretical Analysis

The various embodiments of the MAGS were analyzed using HYSYS code (a chemical process modeling software manufactured by ASPENTECH CORPORATION). The Version 1 MAGS embodiment employed two sequential flashes to remove ethane and residual methane from the liquid stream after the initial liquefaction step and before final stripping of light fluids in the stripping column. The Version 2 MAGS embodiment sent the liquid stream directly to the stripping column immediately after the liquefaction step without any intermediate flash operations.

The theoretical performance obtainable using each of these designs, as predicted by HYSYS, is given in Tables 3 and 4.

TABLE 3 Version 1 Predicted Performance from HYSYS Case 1, Version 1, 180 psi, −60° C., 200 MCF/day Fraction NGL Gas Gas Captured Feed Product Product Product in Gas (C) (A) (B) NGL (%) Stream Specifications Flow Rate. 200 54 113 33 (MCF/day) Pressure, 33 35 103 35 PSIA Temperature, 40 −35 50 −50 ° C. Gas Composition (%) Methane 50.89 0.00 81.78 37.8 0.00 Ethane 21.21 22.64 13.38 53.37 28.29 Propane 13.30 42.64 1.78 6.47 84.96 i-Butane 1.69 6.11 0.06 0.021 95.81 n-Butane 4.33 15.88 0.09 0.031 97.19 i-Pentane 0.82 3.06 0.00 0.001 98.89 n-Pentane 1.08 4.04 0.00 0.001 99.13 n-Hexane 1.50 5.62 0.00 0.00 99.29 CO₂/N₂ 5.18 0.01 2.91 2.31 0.05

TABLE 4 Version 2 Predicted Performance from HYSYS Case 1, Version 2, 180 psi, −60° C., 200 MCF/day Fraction NGL Gas Gas Captured Feed Product Product Product in Gas (C) (A) (B) NGL (%) Stream Specifications Flow Rate. 200 53 113 33 (MCF/day) Pressure, 33 175 103 175 PSIA Temperature, 40 25 50 −20 ° C. Gas Composition (%) Methane 50.89 0.01 81.78 37.57 0.01 Ethane 21.21 22.91 13.38 52.7 28.62 Propane 13.30 42.44 1.78 7.04 84.56 i-Butane 1.69 6.07 0.06 0.3 95.18 n-Butane 4.33 15.81 0.09 0.53 96.76 i-Pentane 0.82 3.06 0.00 0.03 98.89 n-Pentane 1.08 4.04 0.00 0.04 99.13 n-Hexane 1.50 5.63 0.00 0.01 99.46 CO₂/N₂ 5.18 0.03 2.91 1.78 0.15

It can be seen that the Version 2 (Table 4) results are nearly as good as those in Version 1 (Table 3), yet obtained with a significantly simpler system. It would thus be the preferred embodiment of the two. However, both designs are feasible, as are alternatives operating at different temperature and pressure combinations, and utilizing a greater or lesser number of flash tanks, stripping columns, or distillation systems.

Additional Experimental Results

A subscale MAGS unit was built with a capacity of about 14,000 cf/day (14 mcf) of raw natural gas. Both Version 1 and Version 2 configurations were tested. A simulated liquid-rich raw gas feed was created by combining methane, ethane, propane, and butane in the following proportions: methane 50%, ethane 25%, propane 15%, and butane 10%. This gas was then compressed to 180 psi, after which it was refrigerated to −60° C. using an autocascade system, allowing process stream A to vent. Following this step, the liquid stream was either flashed twice and then sent to the stripping column, as per Version 1, or sent directly to the stripping column, as per Version 2.

The results for Version 1 (case 1, no cyclone) are shown in Table 5, while the results for Version 2 (case 2, with cyclonic separator added) are shown in Table 6.

TABLE 5 Case 1 (Version 1, No Cyclone) Experimental Results Experimental Data, Version 1, 130 psia, −55 C, 14 MCF/day Fraction NGL Gas Gas Captured Feed Product Product Product in Gas (C) (A) (B) NGL (%) Stream Specifications Flow Rate, 13.2 2.01 10.07 1.14 (MCF/day) Pressure, 125 20 120 20 PSIA Temperature, 30 −30 −55 −60 ° C. Gas Composition (%) Methane 54.35 1.42 66.57 39.74 0.40 Ethane 18.30 5.97 17.54 46.63 4.96 Propane 16.15 43.03 11.27 11.91 40.48 n-Butane 11.20 49.59 4.62 1.72 67.26

The results from Table 5 show the machine running in Version 1 mode, without a cyclone to remove natural gas liquid droplets in mist form from the exhaust A stream. The composition of the NGL product, which includes both the liquid itself and the head above the liquid, was obtained by mole balance, adjusting Exhaust A flow to obtain methane composition agreement. Gas chromatograph measurements of the liquid portion of the captured NGLs were taken as well. The results were C1=1.84%, C2=6.82%, C3=34.13%, C4=57.22%, agreeing with Table 5.

It can be seen from Table 5, comparing to Table 3, that while the captured liquid was of very high quality, the fraction captured was below theoretical calculations. It was concluded that excessive amounts of liquids were being lost in mist form in the Exhaust A stream. To reduce these losses, a cyclone separator was introduced in addition to testing the machine in Version 2 mode. The results for Version 2 (case 2, with cyclone) are shown in Table 6.

TABLE 6 Case 2 (Version 2, with Cyclonic Separator) Experimental Results Experimental Data, Version 2, 160 psia, −55 C, 14 MCF/day V NGL Gas Gas Captured Feed Product Product Product in Gas (C) (A) (B) NGL (%) Stream Specifications Flow Rate, 13.7 5.15 6.21 2.38 (MCF/day) Pressure, 160 150 155 152 PSIA Temperature, 30 −10 −55 −50 ° C. Gas Composition (%) Methane 54.36 3.57 86.46 80.69 2.46 Ethane 19.01 30.43 11.20 14.64 60.05 Propane 14.77 35.35 1.92 3.69 89.82 n-Butane 11.87 30.66 0.43 0.99 96.94

The results in Table 6 illustrate the machine running in Version 2 mode, with a cyclone added to remove natural gas liquid droplets in mist form from the exhaust A stream. The composition of the NGL product, which includes both the liquid itself and the head above the liquid, was obtained by mole balance, adjusting Exhaust A flow to obtain methane composition agreement. Gas chromatograph measurements of the liquid portion of the captured NGLs were taken as well. The results were C1=3.33%, C2=24.9%, C3=39.9%, C4=31.9%, agreeing with Table 6.

It can be seen that the addition of the cyclonic separator improved NGLs capture, with experimental results achieved closely matching theory. Capture of propane was 89.8% and butane 96.9%, which is highly satisfactory. However, ethane capture in the liquid was potentially higher than desirable, a problem that did not occur in Version 1. This could be remedied in Version 2 by making the stripping column larger. Subsequent to these results, the stripping column was enlarged, leading to the results of the prototype MAGS-0 system shown in Table 2B, which corresponds to a Case 2 system (Version 2, no flash tanks, with cyclonic separator), but with a larger stripping column. Alternatively, the cyclonic separator could be added to the Version 1 system.

Based on these two experimental results, an analysis was made of the projected performance of the Version 1 machine with a cyclonic separator added. The results are shown in Table 7.

TABLE 7 Case 3 (Version 1, with Cyclonic Separator) Projected Performance Experimental Data, Version 1, 130 psia, −55 C, 14 MCF/day Fraction NGL Gas Gas Captured Feed Product Product Product in Gas (C) (A) (B) NGL (%) Stream Specifications Flow Rate, 13.2 3.43 7.63 2.16 (MCF/day) Pressure, 125 20 120 20 PSIA Temperature, 30 −30 −55 −60 ° C. Gas Composition (%) Methane 54.35 1.51 85.84 27.13 0.72 Ethane 18.30 9.41 11.87 55.11 13.35 Propane 16.15 48.55 1.73 15.57 78.05 n-Butane 11.20 40.53 0.55 2.19 93.95

It can be seen from Table 7 that these results approximate those predicted by HYSYS, with the exception that a slightly larger fraction of the propane was lost in exhaust stream B. This is due to liquid propane mist escaping with the gas stream, an effect that was accentuated by the small size of the subscale MAGS experimental unit. Both the quality of the NGLs produced and the NGLs capture fraction are excellent.

System Design Using a Modular Configuration

The present invention may also be configured as a modular system, which may be assembled using modular units (for example, but not limited to, 200 mcf units). These components may be assembled together at the field depending on the particular application, and the requirements of a particular user. Depending on the gas processing needs of a particular site, multiple units may be combined to provide the necessary processing power. Similarly, as production declines or gas gathering lines are added, units can be removed and moved to new production locations.

If necessary, and in some embodiments, a desulfurization subsystem can be added in order to remove any sulfur from the raw gas stream. The desulfurization subsystem could be applied to the gas either upstream of the entire MAGS process, or after the compressor and liquid drop-out (condensers), but before the refrigeration. Several sulfur treatment and removal methods are possible. Dry sorbents may be used to capture sulfur in the feed gas. Calcium oxide, magnesium oxide, and sodium carbonate are example dry sorbents that are capable of trapping sulfur gases in solid form (as sulfates or sulfites, depending on the relative oxidation conditions). A fine sorbent can be injected into the feed gas, with resulting sulfur containing solids then collected. In other embodiments, sulfur may also be removed by using a wet scrubber subsystem. Wet scrubbers can be configured in venturi, packed-column, or tray-type systems in which the feed gas is contacted with a scrubbing solution or slurry. The resulting scrubber solution or slurry must then be disposed.

Various Use Cases of the Present Invention

Several alternative use cases of the present invention are now presented. These use cases are illustrative of the possible applications of the present invention and are not meant to be exhaustive or limiting.

FIG. 8 illustrates an example of a use case 800 of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is used for electricity production. As shown in FIG. 8, oil and associated gas (which is normally flared) is produced from well 802, which may be a well fracked with hydraulic fracturing as practiced in North Dakota. The flare gas is taken to MAGS unit 804, where it is separated into NGLs 806 and lean methane, which can be utilized to power an engine/electric generator combo 808. The produced electricity may be used on-site to power various oil field equipment 812, or transmitted to the grid 814.

FIG. 9 illustrates another example of a use case 900 of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is compressed into CNG, and used as vehicle fuel or transported to a remote location to power remote gensets. As shown in FIG. 9, oil and associated gas (which is normally flared) is produced from well 902. The flare gas is taken to MAGS unit 904, where it is separated into NGLs 906 and lean methane, which can be compressed with compressor 908 to fill CNG storage cylinders 910. The stored CNG may be off-loaded directly to CNG vehicle tanks 912 for use in vehicle transportation, or loaded into CNG transport tankers 914, which may be transported to remote locations (for example, to fuel CNG fleets or provide CNG to drilling rigs).

FIG. 10 illustrates yet another example of a use case 1000 of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is converted to liquid fuels for easy transport to a remote site. As shown in FIG. 10, oil and associated gas (which is normally flared) is produced from well 1002. The flare gas is taken to MAGS unit 1004, where it is separated into NGLs 1006 and lean methane. The methane may then be upgraded to various liquid fuels 1010, including methanol, ethanol, propanol, butanol, etc. via reformer/synthesis unit 1008. The reformer/synthesis unit 1008 may include a reformer which produces syngas, followed by a synthetic fuel production subsystem (see, for example, U.S. Pat. No. 8,450,536, issued on May 28, 2013, to Emily B. White, et al., and entitled “Methods of Higher Alcohol Synthesis,” which shows one exemplary system and method for converting synthesis gas, which can be derived from natural gas by steam reforming (see below for discussion on how to produce synthesis gas on-site), into liquid fuels comprising methanol, ethanol, propanol, and butanol, and higher hydrocarbons, the entirety of which is hereby incorporated by reference herein). Other liquid fuels and butanol synthesis pathways are within the scope of the present invention. In one preferred embodiment, the lean methane A-gas stream is converted into methanol, with the methanol delivered by truck to sites of CO₂ or CH₄ availability, and then the methanol is upgraded to a higher alcohol fuel, such as butanol (see, for example, U.S. Patent Application Publication No. 2013/0118063 to John Henri, et al., entitled “Synthesis of High Caloric Fuels and Chemicals,” which shows several exemplary systems and methods for converting synthesis gas, methanol, or acetic acid into butanol, the entirety of which is hereby incorporated by reference herein). The liquid fuels 1010 may be collected on-site in storage containers and transported from the oil site using truck-trailers in the same manner as the oil is transported from the site.

FIG. 11 illustrates yet another example of a use case 1100 of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is used to create an artificial gas cap or for enhanced oil recovery. As shown in FIG. 11, oil and associated gas (which is normally flared) is produced from well 1102. The flare gas is taken to MAGS unit 1104, where it is separated into NGLs 1106 and lean methane. The methane may be compressed via compressor 1108 and re-injected into the same or nearby petroleum formation 1112 via an injection well 1110 to create or re-establish an artificial gas cap.

FIG. 12 illustrates yet another example of a use case 1200 of the mobile alkane gas separator (MAGS) in which the NGLs are transported and the CH₄ stream is converted to CO₂ and H₂, which are then used for enhanced oil recovery and electricity production, respectively. As shown in FIG. 12, oil and associated gas (which is normally flared) is produced from well 1202. The flare gas is taken to MAGS unit 1204, where it is separated into NGLs 1206 and lean methane. The methane may be used to power a Portable Enhanced Recovery Technology (PERT) system 1208, which steam reforms the natural gas into CO₂ and H₂ (see, for example, U.S. Pat. No. 7,918,906, issued on Apr. 5, 2011, to Robert M. Zubrin, et al., and entitled “Compact Natural Gas Steam Reformer with Linear Countercurrent Heat Exchanger,” which shows one exemplary system and method for reforming natural gas into CO₂ and H₂ using a portable apparatus brought to an oil field for use in EOR, the entirety of which is hereby incorporated by reference herein). The same PERT system may also be used to manufacture synthesis gas (CO and H₂) needed for liquid fuel production on the site of a flare well as described in relation to FIG. 10. The CO₂ is compressed via compressor 1210 and injected into the same or nearby petroleum formation 1214 via injection well 1212 for enhanced oil recovery (CO₂-EOR). The H₂ may be utilized for electricity production 1216, or for other purposes, such as hydrogenation of the produced oil. The CO₂ may also be used for waterless CO₂-fracking to re-fracture the same well or to fracture a nearby well.

Potential Macro-Environmental and Macroeconomic Impact

Previously, the impact of the technology on a single user was discussed, to show that it would be highly profitable. This is the key to the propagation of the technology to a large number of fields. In this section, the macro-environmental and macroeconomic effect of the technology is discussed once it has been put into broad use, showing that it could have a major impact in both increasing NGLs utilization, meeting expanded electricity needs of drillers, and reducing carbon emissions.

In North Dakota alone, 190,000 mcf per day of natural gas is being flared, approximately 60% (114,000 mcf/day) of which is coming from wells producing 200 mcf per day or more each. That is sufficient market, in North Dakota alone, for 570 MAGS-200 units. If there was 50% market penetration (57,000 mcf/day) flaring reduced by MAGS-200, then almost 3,000 metric tons of wasted CO₂-equivalent emissions per day would be avoided. This translates to ˜1.1 million tons of CO₂-equivalent (1.1 Mt) emissions avoided per year. Meanwhile, an economic gain of $208 million to $513 million per year is expected from the sale of the NGLs and utilization of the methane for electricity production in North Dakota alone.

The United States oil and gas industry annually flared approximately 7.1 billion cubic meters (bcm), or 250 billion cubic feet (bcf) in 2011. If just 15% of this flaring was avoided through the use of MAGS, then ˜2.0 Mt of CO₂-equivalents would be avoided per year in the United States. These represent not just significant environmental damage, but significant economic opportunity in the United States. This would represent an economic opportunity of $378 million to $932 million per year.

Meanwhile, it is estimated that Canada flares 2.4 billion m³ per year (Source: Global Gas Flaring Reduction Partnership, Estimated Flared Volumes from Satellite Data, 2007-2011, 2013.) If just 45% of this flaring was avoided through the use of MAGS in Canada (Canada's flare sites are more concentrated), then an additional ˜2.0 Mt of CO₂-equivalents would be avoided per year in Canada. These represent not just significant environmental damage, but significant economic opportunity in Canada as well. With the use of the MAGS, both Canada and the United States could achieve greater energy independent in North America by maximizing NGLs capture and simultaneously reducing greenhouse gas emissions from the North American oil and gas sector.

Long-Felt, Unsolved Need for Cost-Effective on-Site Gas Separation—North Dakota Case Study

As stated by a recent study conducted for the U.S. Department of Energy, around 30% of North Dakota's produced associated gas is flared, nearly 190 million cubic feet per day (190,000 mcf/day). (Source: Wocken, C. A.; Stevens, B. G.; Almlie, J. C.; Schlasner, S. M., End-Use Technology Study—An Assessment of Alternative Uses for Associated Gas, National Energy Technology Laboratory, Pittsburgh, Pa., April 2013, incorporated by reference in its entirety herein.) This study demonstrates the long-felt and unsolved need for mobile technology to address this issue. This study also shows that no existing technology can simultaneously produce a lean methane stream as well as remove NGLs from flare gas, a key innovation of the present patent application. This discussion is merely illustrative and exemplary, and is not intended to limit the scope of the present invention or its application or uses.

CONCLUSION

The inventors have completed the construction of the full-scale, MAGS-200 portable unit, have moved the unit from the company facilities to the field, and are beginning field testing in Colorado.

While the methods disclosed herein have been described and shown with reference to particular operations performed in a particular order, it will be understood that these operations may be combined, sub-divided, or re-ordered to form equivalent methods without departing from the teachings of the present invention. Accordingly, unless specifically indicated herein, the order and grouping of the operations is not a limitation of the present invention.

While the present invention has been particularly shown and described with reference to embodiments thereof, it will be understood by those skilled in the art that various other changes in the form and details may be made without departing from the spirit and scope of the present invention. 

What is claimed is:
 1. An apparatus for separating methane and natural gas liquids from a raw natural gas stream, comprising: a chassis adapted to hold the system for field deployment; one or more compressors for compressing the raw natural gas stream; one or more dehydrators for removing water from the compressed natural gas stream; a refrigerator having one or more stages for lowering a temperature of the dehydrated, compressed natural gas stream; and a separation subsystem system adapted to separate the dehydrated, compressed, and refrigerated natural gas stream into three product streams consisting essentially of a methane stream that is at least 80% methane, an ethane-rich stream that is high in ethane content, and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C.
 2. The apparatus of claim 1, further comprising: a power generator tuned to run on ethane-rich gas adapted to provide electricity to run the compressor, the refrigerator, and other system components, utilizing the ethane-rich stream as its energy source.
 3. The apparatus of claim 1, wherein the methane stream is at least 85% methane.
 4. The apparatus of claim 1, wherein the ethane-rich stream is at least 30% ethane.
 5. The apparatus of claim 1, wherein the natural gas liquids stream has a vapor pressure of no more than 200 psia at 38° C.
 6. The apparatus of claim 1, wherein the natural gas liquids stream meets a Y grade standard.
 7. The apparatus of claim 1, wherein the one or more compressors compress the raw natural gas stream to a pressure range of 100 psig to 500 psig.
 8. The apparatus of claim 1, wherein the chassis is mounted on a trailer having one or more wheels.
 9. The apparatus of claim 1, wherein the refrigerator further comprises: a high-stage refrigeration loop having at least one heat exchanger for lowering a temperature of the dehydrated, compressed natural gas stream; and a low-stage refrigeration loop having at least one heat exchanger for further lowering the temperature of the dehydrated, compressed natural gas stream.
 10. The apparatus of claim 9, wherein the low-stage refrigeration loop is an autocascade loop having mixed refrigerants.
 11. The apparatus of claim 10, wherein the mixed refrigerants are hydrocarbons.
 12. The apparatus of claim 1, wherein the refrigerator cools the natural gas stream to a temperature range of −40° C. to −70° C. sufficient to achieve a NGLs separation in a single separation column.
 13. The apparatus of claim 1, wherein one or more of the dehydrators employ a desiccant bed.
 14. The apparatus of claim 13, wherein two desiccant beds are employed in alternation, and wherein heat required to dry the two desiccant beds is derived from waste heat from a power generator that drives the compressors and the refrigerator.
 15. The apparatus of claim 1, wherein the separation subsystem comprises a stripping column.
 16. The apparatus of claim 1, wherein the separation subsystem comprises a distillation column.
 17. The apparatus of claim 1, wherein the separation subsystem comprises one or more flash tanks.
 18. The apparatus of claim 1, wherein the separation subsystem comprises one or more cyclones to separate liquids from gasses.
 19. A method for separating useable gas streams from a raw natural gas stream, comprising: a compression step for compressing the raw natural gas stream utilizing a compressor to a pressure range of 100 to 500 psig; a dehydration step for removing water from the compressed natural gas stream utilizing one or more condensers and one or more desiccant beds; a refrigeration step for reducing a temperature of the dehydrated, compressed natural gas stream utilizing a refrigerator having one or more stages to a temperature range of −40° C. to −70° C.; and a separation step for separating the dehydrated, compressed, and refrigerated natural gas stream into three product streams consisting essentially of a methane stream of at least 80% methane, an ethane-rich stream that is high in ethane content, and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C.
 20. The method of claim 19, further comprising: a generation step for generating power utilizing a power generator tuned to run on ethane-rich gas to provide electricity to run the compressor and the refrigerator, utilizing the ethane-rich stream as its energy source.
 21. The method of claim 19, further comprising: a transportation step for bringing a mobile alkane gas separator system to an oil field site that is flaring associated gas; and a deployment step for connecting said mobile alkane gas separator system to a raw associated natural gas source.
 22. The method of claim 19, wherein the methane stream is at least 85% methane.
 23. The method of claim 19, further comprising: compressing the methane stream to produce compressed natural gas (CNG).
 24. The method of claim 19, further comprising: converting the methane stream into methanol for transport; transporting the methanol to a remote site; and upgrading the methanol to higher alcohols at the remote site.
 25. The method of claim 19, further comprising: utilizing the methane stream for enhanced oil recovery.
 26. The method of claim 19, wherein the refrigeration step utilizes an autocascade refrigerator having mixed hydrocarbon refrigerants.
 27. The method of claim 19, wherein the refrigeration step cools the natural gas stream to at temperature range of −50° C. to −60° C.
 28. The method of claim 19, wherein two desiccant beds are employed in alternation, and wherein heat required to dry the two desiccant beds is derived from waste heat from a power generator.
 29. A system for separating a raw natural gas stream into a methane-enriched stream, an ethane-enriched stream, and a natural gas liquids stream, comprising: one or more compressors for compressing the raw natural gas stream; a dehydrator subsystem for removing water from the compressed natural gas stream; a refrigeration subsystem for lowering a temperature of the dehydrated, compressed natural gas stream to a low temperature, comprising a high-stage cascade refrigeration loop and a low-stage autocascade refrigeration loop having mixed hydrocarbon refrigerants; a separation subsystem adapted to separate the dehydrated, compressed, and refrigerated natural gas stream into three product streams consisting essentially of a methane stream that is at least 80% methane, an ethane-rich stream that is high in ethane content, and a natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C.; and a power generation subsystem tuned to run on ethane-rich gas adapted to provide electricity to run the compressor, the refrigerator, and other system components, utilizing the ethane-rich stream as its energy source.
 30. The system of claim 29, wherein the dehydration subsystem employs two desiccant beds in alternation, wherein heat required to dry the two desiccant beds is derived from waste heat from the power generator.
 31. The system of claim 29, wherein the natural gas liquids stream has a vapor pressure of no more than 200 psia at 38° C.
 32. The system of claim 29, further comprising a chassis for holding system components for field deployment, said chassis mountable to a trailer having one or more wheels for transportation to a field location.
 33. The system of claim 29, wherein the refrigeration subsystem comprises: a high-stage refrigeration loop having at least one heat exchanger for lowering a temperature of the dehydrated, compressed natural gas stream; and a low-stage refrigeration loop having at least one heat exchanger for further lowering the temperature of the dehydrated, compressed natural gas stream, wherein the low-stage refrigeration loop is an autocascade loop having mixed refrigerants.
 34. The system of claim 29, wherein the refrigeration subsystem cools the natural gas stream to a temperature range of −40° C. to −70° C. sufficient to achieve a NGLs separation in a single separation column.
 35. A natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C. produced from a raw natural gas stream at a wellhead, said natural gas liquids stream made by a process comprising the steps of: compressing the raw natural gas stream utilizing a compressor to a pressure range of 100 to 300 psig; removing water from the compressed natural gas stream utilizing one or more condensers and one or more desiccant beds to achieve an aqueous dew point below −70° C.; reducing a temperature of the dehydrated, compressed natural gas stream utilizing a refrigerator having one or more stages to a temperature range of −40° C. to −70° C.; and separating the dehydrated, compressed, and refrigerated natural gas stream into said natural gas liquids stream having a vapor pressure of no more than 250 psia at 38° C. and two additional product streams consisting essentially of a methane stream of at least 80% methane and an ethane-rich stream that is high in ethane content. 